Fluidizable vanadium catalyst for oxidative dehydrogenation of alkanes to olefins in a gas phase oxygen free environment

ABSTRACT

Fluidizable catalysts for the gas phase oxygen-free oxidative dehydrogenation of alkanes, such as propane, to corresponding olefins, such as propylene. The catalysts comprise 5-20% by weight per total catalyst weight of one or more vanadium oxides (VO x ), such as V 2 O 5 . The dehydrogenation catalysts are disposed on an alumina support that is modified with calcium oxide to influence characteristics of lattice oxygen at the catalyst surface. Various methods of preparing and characterizing the catalyst as well as methods for the gas phase oxygen free oxidative dehydrogenation of alkanes, such as propane, to corresponding olefins, such as propylene, with improved alkane conversion and olefin product selectivity are also disclosed.

BACKGROUND OF THE INVENTION Technical Field

The present disclosure relates to fluidizable vanadium basedVO_(x)/CaO-γ-Al₂O₃ catalysts and dehydrogenation processes using thecatalysts for the oxidative dehydrogenation of alkanes, such as propaneto propylene, in the absence of gas phase oxygen.

Description of the Related Art

The “background” description provided herein is for the purpose ofgenerally presenting the context of the disclosure. Work of thepresently named inventors, to the extent it is described in thisbackground section, as well as aspects of the description which may nototherwise qualify as prior art at the time of filing, are neitherexpressly or impliedly admitted as prior art against the presentinvention.

Propylene is an important precursor and feedstock in the chemicalindustry used to produce many different valuable products. Approximatelytwo thirds of the propylene produced worldwide is consumed in theproduction of thermoplastic polypropylene which is commonly used in thefabrication of household appliances, plastic films and many otherapplications. With the increasing world population and the improvingquality of human life, worldwide propylene demand/sales has reached overninety billion dollars [“Market Study: Propylene, Ceresana Research,February 2011.” ceresana.com. Retrieved 2011-02-13]. Conventionally,propylene has been produced from petroleum refining and olefin crackingprocesses. In order to meet the ever increasing demands for petroleumfuel and olefins (propylene, ethylene, etc.) there is a growing need todevelop alternative propylene production technology. In this regard,propylene from propane, available both in natural gas and refinery offgases, is considered an attractive technology. The abundant availabilityof propane in different parts of the world including the United Statesand the Middle Eastern Region can make these intentional propyleneproductions largely sustainable as compared to refinery and olefincracker processes [Ashford's Dictionary of Industrial Chemicals, ThirdEdition, 2011, ISBN 978-0-9522674-3-0, pages 7766-9].

At present, there are three major commercial processes available in theproduction of propylene, including steam cracking, fluid catalyticcracking (FCC) and catalytic dehydrogenation. Steam cracking processesconsume a large amount of heat energy, which accounts for 70% of theoverall production cost. Coke formation during the cracking of heavyhydrocarbon molecules is an additional drawback of the steam crackingprocess. This coke formation causes severe operational constraints onthe process, especially fouling, which requires frequent plant shutdowns for cleaning. In contrast, in the catalytic cracking (FCC) processthe coke generation is deliberate. The formed coke is combusted in thecatalyst regenerator producing heat energy that is supplied back to thecatalytic cracking unit using the catalyst as an energy carrier. Thisenergy is essential for the FCC reactor to carry out the endothermiccracking reactions. In the FCC process the propylene is obtained as abyproduct, in addition to the lighter gasoline and other fuels. Theyield of propylene can be increased by manipulating the FCC operatingconditions and using catalysts as additives. Recent research shows thatthe FCC catalysts can improve the propylene yield by about 4.5% to 10%.However, the propylene production cost in the FCC process is still highdue to the energy required by the endothermic cracking reactions. Thishigh energy demand and the continuous catalyst regeneration, make theFCC process a capital intensive one. Consequently, the building of a FCCsystem for the sole aim of producing propylene is not economicallyviable. The third available technology, catalytic dehydrogenation, alsosuffers from the problems of coke formation and high energy requirementsas a result of the endothermic nature of the reaction [S. A. Al-Ghamdi(2013) oxygen-free propane oxidative dehydrogenation over vanadium oxidecatalysts: reactivity and kinetic modeling. Ph.D dissertationmonograph.—incorporated herein by reference in its entirety].

Contrary to the above discussed commercial processes, the oxidativedehydrogenation (ODH) of propane to propylene is more attractive due toits low operational cost and minimal environmental impacts. Compared tothe present commercial technologies, oxidative dehydrogenation reducescosts, saves energy, and lowers greenhouse gas emissions. The mostimportant advantage is the exothermic nature of the reaction, whichrequires no additional energy to accelerate the reaction. The formationof water as a byproduct of the oxidative dehydrogenation makes itpossible to avoid thermodynamic constraints that are observed in thenon-oxidative routes. The activity of the catalyst is also stable for alonger number of cycles due to the minimal coke deposition at thecatalyst surface [E. Heracleous, M. Machli, A. A. Lemonidou, I. A.Vasalos, Oxidative dehydrogenation of dehydrogenation of ethane andpropane over vanadia and molybdena supported catalysts, J. Mol. Catal.A: Chem. 232 (2005) 29-39.—incorporated herein by reference in itsentirety]. It is has been thought that high propylene yield can beobtained through the ODH of propane with the successful development ofefficient catalysts. In the ODH process the operation and maintenancecosts are relatively low as a result of the low operating temperature.The use of a furnace and the need for decoking shutdowns are also notessential parts of the oxidative dehydrogenation process. This allaccounts for comparatively small capital investment for implementationof ODH processes while still providing appreciable operationalefficiencies.

The selection of a reactor is very important for the commercial scaleapplication of oxidative dehydrogenation technology [L. Chalakov, L. K.Rihko-Struckmann, B. Munder, K. Sundmacher, Oxidative dehydrogenation ofethane in an electrochemical packed-bed membrane reactor: Model andexperimental validation, Chem. Eng. J. 145 (2009) 385-392.—incorporatedherein by reference in its entirety]. Fixed bed reactors are simple butdifficult to maintain at isothermal conditions which can causeinterference with the performance of the reactor and lead to catalystdamage and deactivation. There are numerous advantages to fluidized bedreactors over the conventional fixed reactor systems. These includecontrolled operation conditions at constant temperature, which assist incircumventing the issues associated with hot spots in fixed bedreactors. The absence of mass transfer limitations and uniform residencetime distributions (RTDs) are also merits of fluidized bed reactors.Moreover, the ability to transport reduced catalytic species from theoxidative dehydrogenation unit to a regeneration unit is another meritof fluidized bed reactors that poses the opportunity for periodiccatalyst re-oxidation. This enables twin reactor set ups, with one foroxidative dehydrogenation and the other for regeneration of thecatalyst, which is important for commercial scale production [S. A.Al-Ghamdi, M. Volpe, M. M. Hossain, H. I. de Lasa, VO_(x)/c-Al₂O₃catalyst for oxidative dehydrogenation of ethane to ethylene: desorptionkinetics and catalytic activity. Appl. Catal. A: Gen. 450 (2013)120-130; and A. W. H. Elbadawi, M. S. Ba-Shammakh, S. A. Al-Ghamdi, S.A. Razzak, M. M. Hossain, Reduction kinetics and catalytic activity ofVO_(x)/γ-Al₂O₃—ZrO₂ for gas phase oxygen free ODH of ethane, Chem. Eng.J. 284 (2016) 448-457; and I. A. Bakare, M. Shamseldin, S. A. Razzak, S.A. Al-Ghamdi, M. M. Hossain (2015), H. I. de Lasa, Fluidized bed ODH ofethane to ethylene over VO_(x)—MoO_(x)/γ-Al₂O₃ catalyst: Desorptionkinetics and catalytic activity, Chem. Eng. J. 278 (2015) 207-216.—eachincorporated herein by reference in its entirety].

In addition to the proper reactor selection, the development of suitablecatalysts is another key aspect for the commercial implementation of ODHprocesses. High yield of propylene can be obtained by employing anefficient catalyst. It has been previously reported in the literature onoxidative dehydrogenation that vanadium based catalysts offer thehighest alkane conversion and alkene selectivity from ethane and otherlighter hydrocarbons in the ODH reaction [M. M. Bhasin, Is true ethaneoxydehydrogenation feasible, Top. Catal. 4 (2003) 145-149; and L. Capek,R. Bulanek, J. Adam, L. Smolakova, H. Sheng-Yang, P. Cicmanec, Oxidativedehydrogenation of ethane over vanadium-based hexagonal mesoporoussilica catalysts, Catal. Today 141 (2009) 282-287; and L. Capek, J.Adam, T. Grygar, R. Bulanek, L. Vradman, G. Kosova-Kucerova, P.Cicmanec, P. Knotek, Oxidative dehydrogenation of ethane over vanadiumsupported on mesoporous materials of M41S family, Appl. Catal. A: Gen.342 (2008) 99-106; and F. Klose, T. Wolff, H. Lorenz, A.Seidelmorgenstern, Y. Suchorski, M. Piorkowska, H. Weiss, Active specieson γ-alumina-supported vanadia catalysts: Nature and reducibility J.Catal. 247 (2007) 176-193; and A. Klisinska, S. Loridant, B. Grzybowska,J. Stoch, I. Gressel, Effect of additives on properties of V₂O₅/SiO₂ andV₂O₅/MgO catalysts II. Structure and physicochemical properties of thecatalysts and their correlations with oxidative dehydrogenation ofpropane and ethane, Appl. Catal. A: Gen. 309 (2006) 17-27; and B.Grzybowska, A. Klisinska, K. Samson, I. Gressel, Effect of additives onproperties of V₂O₅/SiO₂ and V₂O₅/MgO catalysts: I. Oxidativedehydrogenation of propane and ethane Appl. Catal. A: Gen. 309 (2006)10-16; and M. V. Martinez-Huerta, X. Gao, H. Tian, I. E. Wachs, J. L. G.Fierro, M. A. Banares, Oxidative dehydrogenation of ethane to ethyleneover alumina-supported vanadium oxide catalysts: Relationship betweenmolecular structures and chemical reactivity, Catal. Today 4 (2006)279-287; and R. Grabowski, J. Sloczynski, Kinetics of oxidativedehydrogenation of propane and ethane on VOx/SiO₂ pure and withpotassium additive, Chem. Eng. Process. 44 (2005) 1082-1093; and E. P.Reddy, R. S. Varma, Preparation, characterization, and activity ofAl₂O₃-supported V₂O₅ catalysts, J. Catal. 221 (2004) 93-101; and F.Bozon-Verduraz, D. I. Enache, E. Bordes, A. Ensuque, Vanadium oxidecatalysts supported on titania and zirconia: II. Selective oxidation ofethane to acetic acid and ethylene, Appl. Catal. A: Gen. 278 (2004)103-110; and D. I. Enache, E. Bordes, A. Ensuque, F. Bozon-Verduraz,Vanadium oxide catalysts supported on zirconia and titania: I.Preparation and characterization. Appl. Catal. A: Gen. 278 (2004)93-102; and P. Concepcion, M. T. Navarro, J. M. Lopez-Nieto, T. Blasco,B. Panzacchi, F. Rey, Vanadium oxide supported on mesoporous Al₂O₃:Preparation, characterization and reactivity, Catal. Today 96 (2004)179-186; and Z. Zhao, Y. Yamada, A. Ueda, H. Sakurai, T. Kobayashi, Theroles of redox and acid-base properties of silica-supported vanadiacatalysts in the selective oxidation of ethane, Catal. Today 95 (2004)163-171; and G. Busca, M. Panizza, C. Resini, F. Raccoli, R. Catani, S.Rossini, Oxidation of ethane over vanadia-alumina-based catalysts:co-feed and redox experiments, Chem. Eng. J. 93 (2003) 181-189; and A.T. Bell, E. Iglesia, M. D. Argyle, K. Chen, Ethane OxidativeDehydrogenation Pathways on Vanadium Oxide Catalysts, J. Phys. Chem. B106 (2002) 5421-5427; and H. I. de Lasa, M. Volpe, G. Tonetto, Butanedehydrogenation on vanadium supported catalysts under oxygen freeatmosphere. Appl. Catal. A: Gen. 272 (2004) 69-78.—each incorporatedherein by reference in its entirety]. This is attributed to vanadiumcatalysts providing lattice oxygen for the dehydrogenation of alkanes[A. T. Bell, A. Dinse, R. Schomacker, The role of lattice oxygen in theoxidative dehydrogenation of ethane on alumina-supported vanadium oxide,Phys. Chem. Chem. Phys. 29 (2009) 6119-6124; and E. A. Mamedov, V. C.Corberfin, Oxidative dehydrogenation of lower alkanes on vanadiumoxide-based catalysts. The present state of the art and outlooks. Appl.Catal. A: Gen. 127 (1995) 1-40.—each incorporated herein by reference inits entirety].

The reactions involved in the oxidation of propane include the desiredpropane oxidative dehydrogenation to propylene as well as the combustionof propane and produced propylene to carbon (IV) oxides and carbon (II)oxide. High selectivity for propylene is thus only favorable at lowpropane conversions due to the lower reactivity of propane when comparedto propylene. Thus, there is a need to design catalysts that willprovide lattice oxygen that can selectively produce propylene from theODH of propane while minimizing or preventing the primary and secondarycombustion of propane and propylene respectively to carbon oxides [S. A.Al-Ghamdi, H. I. de Lasa, Propylene production via propane oxidativedehydrogenation over VOx/γAl₂O₃ catalyst. Fuel 128 (2014)120-140.—incorporated herein by reference in its entirety].

The performance of supported vanadium oxide in oxidative dehydrogenationreactions is a function of the redox properties and morphology ofsurface species of VO_(x) and the acid-base character of the VO_(x)catalyst and its support [A. Khodakov, B. Olthof, A. T. Bell, E.Iglesia, Structure and catalytic properties of supported vanadiumoxides: support effects on oxidative dehydrogenation reactions. J.Catal. 181 (1999) 205-216; and M. V. Martinez-Huerta, X. Gao, H. Tian,I. E. Wachs, J. L. G. Fierro, M. A. Banares, Oxidative dehydrogenationof ethane to ethylene over alumina-supported vanadium oxide catalysts:relationship between molecular structures and chemical reactivity.Catal. Today 118 (2006) 279-287; and M. A. Banares, Supported metaloxide and other catalysts for ethane conversion: a review, Catal. Today51 (1999) 319-348; and I. E. Wachs, B. M. Weckhuysen, Structure andreactivity of surface vanadium oxide species on oxide supports, Appl.Catal. A: Gen. 157 (1997) 67-90; and D. I. Enache, E. Bordes-Richard, F.Bozon-Verduraz, A. Ensuque, Vanadium oxide catalysts supported onzirconia and titania I. Preparation and characterization, Appl. Catal.A: Gen. 278 (2004) 93-102.—each incorporated herein by reference in itsentirety]. Vanadium catalyst activity and selectivity is a function ofthe structure of VO_(x) surface species. The surface density of VO_(x)increases with vanadium loading, which is lowest for monovanadateisolated VO_(x) species and highest for monolayer coverage species.Catalyst activity and reducibility increases as the surface density ofVO_(x) increases, while selectivity decreases as the surface density ofVO_(x) increases [J. M. Lopez-Nieto, J. Soler, P. Concepcion, J.Herguido, M. Menendez, J. Santamaria, Oxidative Dehydrogenation ofAlkanes over V-based Catalysts: Influence of Redox Properties onCatalytic Performance, J. Catal. 185 (1999) 324-332; and K. Chen, A. T.Bell, E. Iglesia, The relationship between the electronic and redoxproperties of dispersed metal oxides and their turnover rates inoxidative dehydrogenation reactions, J. Catal. 209 (2002) 35-42; and F.Roozeboom, M. C. Mittelmeijer-Hazeleger, J. A. Moulijn, J. Medema, V. H.J. Beer De, P. J. Gellings, Vanadium oxide monolayer catalysts. 3. ARaman spectroscopic and temperature programmed reduction study ofmonolayer and crystal type vanadia on various supports, J. Phys. Chem.84 (1980) 2783-2791; and J. M. Lopez-Nieto, The selective oxidativeactivation of light alkanes from supported vanadia to multicomponentbulk V-containing catalysts. Top Catalysis 41 (2006) 3-15; and G.Che-Galicia, R. Quintana-Solorzano, R. S. Ruiz-Martinez, J. S. Valente,C. O. C. Araiza, Kinetic modeling of the oxidative dehydrogenation ofethane to ethylene over a MoVTeNbO catalytic system, Chem. Eng. J. 252(2014) 75-88—each incorporated herein by reference in its entirety].Adjustments of the coordination and environment of the species of VO_(x)can influence its catalytic behavior. The acid-base character of VO_(x)catalyst supports has been shown in past research to have an influenceon propylene selectivity in the ODH reaction. Propane adsorption andpropylene desorption are functions of the acid-base properties of thesupport. Adsorption of the basic reactant and desorption of the acidicproduct are functions of the acidity of the catalyst. The acidity of thecatalyst determines the protection of these chemical species fromfurther oxidizing to carbon oxides [H. Kung, P. M. Michalakos, M. C.Kung, I. Jahan, Selectivity patterns in alkane oxidation overMg₃(VO₄)₂—MgO, Mg₂V₂O₇, and (VO)₂P₂O₇. J. Catal. 140 (1993)226-242.—incorporated herein by reference in its entirety]. The acidiccharacter of alkanes and their corresponding olefins diminishes withincreased carbon numbers and degree of molecule saturation [J.Santander, E. López, A. Diez, M. Dennehy, M. Pedernera, G. Tonetto,Ni—Nb mixed oxides: One-pot synthesis and catalytic activity foroxidative dehydrogenation of ethane, Chem. Eng. J. 255 (2014) 185-194;and J. P. Bortolozzi, T. Weiss, L. B. Gutierrez, M. A. Ulla, Comparisonof Ni and Ni—Ce/Al₂O₃ catalysts in granulated and structured forms:Their possible use in the oxidative dehydrogenation of ethane reaction,Chem. Eng. J. 246 (2014) 343-352.—each incorporated herein by referencein its entirety].

There are typically strong interactions between the support (carrier)and the active site (VO_(x)). Gamma aluminum oxide (γ-Al₂O₃) is notinert towards VO_(x). Its interactions towards a VO_(x) phase are notweak, and hence can result in a very high dispersion of V₂O₅ on itssurface. High vanadium loading can be achieved on γ-Al₂O₃, but may havelower surface areas unlike calcium oxide (CaO) which has a highersurface area. While the use of CaO may improve the resulting catalyst'ssuperficial area and also provide the desired moderate level of acidity,it may also minimize propylene and propane combustion. Hence, thesynthesis of mixed γ-Al₂O₃/CaO supports is an interesting route toexamine to achieve catalyst samples with high dispersion of the surfaceactive species and a surface area that is higher than that of γ-Al₂O₃[N. E. Quaranta, J. Soria, V. Cortés Corbéran, J. L. G. Fierro,Selective Oxidation of Ethanol to Acetaldehyde on V₂O₅/TiO₂/SiO₂Catalysts J. Catal. 171 (1997) 1-13.—incorporated herein by reference inits entirety].

Ahmed, et al. (U.S. Pat. No. 8,609,568B2) describes a catalyst for theODH of propane to propylene. The disclosed catalyst obtained 5.62%,69.68%, and 8.06% as yield of propylene, selectivity of propylene andpropane conversion respectively in a fixed bed reactor over a 12.5%Ni/VMCM41 catalyst at 400° C. Brophy, et al. (EP1546072A2) describescatalytic oxidative dehydrogenation and microchannel reactors forcatalytic oxidative dehydrogenation. The disclosed catalyst obtained23%, 34.9% and 65.9% as yield of propylene, selectivity of propylene andpropane conversion respectively over a Mg—Mo—V—O catalyst at 583° C.Ahmed, et al. (U.S. Pat. No. 8,623,781A1) describes the oxidativedehydrogenation of propane. The disclosed catalyst obtained 3.43%, 100%,and 3.43% as yield of propylene, selectivity of propylene and propaneconversion respectively in a fixed bed reactor over aMo_(0.5)V_(0.5)O_(5.5) catalyst at 450° C.

In view of the forgoing, one aspect of the present invention is toprovide fluidizable dehydrogenation catalysts comprising vanadium oxidecatalytic species using a mixed γ-Al₂O₃—CaO with different CaO toγ-Al₂O₃ ratios, such as 1:1 or 1:4, as support material. Thephysicochemical characterization of these catalysts offers anexamination of the monovanadate and polyvanadate catalytic VO_(x)surface species, the catalyst's oxygen carrying capacity, level ofacidity and active site metal-support interactions. A further aim of thepresent disclosure is to provide methods for producing theseVO_(x)/CaO-γ-Al₂O₃ catalysts. An additional aim of the presentdisclosure is to provide methods for the oxidative dehydrogenation of analkane, such as propane, to a corresponding olefin, such as propylene,employing the lattice oxygen of these VO_(x)/CaO-γ-Al₂O₃ catalysts.These methods may be performed in a gas phase oxygen free environmentunder fluidized bed reaction conditions while accomplishing high alkaneconversion and high olefin product selectivity over CO_(x) combustionproducts by appropriate control of the lattice oxygen of the catalysts.

BRIEF SUMMARY OF THE INVENTION

According to a first aspect, the present disclosure relates to adehydrogenation catalyst comprising i) a support material comprisingalumina modified by calcium oxide and ii) a catalytic materialcomprising one or more vanadium oxides disposed on the support material,wherein the dehydrogenation catalyst comprises 5-20% of the one or morevanadium oxides by weight relative to the total weight of thedehydrogenation catalyst.

In one embodiment, the weight ratio of calcium oxide to alumina is inthe range of 1:0.2 to 1:8.

In one embodiment, the weight ratio of calcium oxide to alumina is 1:1.

In one embodiment, the one or more vanadium oxides form an amorphousphase on the surface of the support material.

In one embodiment, the one or more vanadium oxides are at least oneselected from the group consisting of V₂O₅, VO₂, and V₂O₃.

In one embodiment, the dehydrogenation catalyst comprises at least 50%of V₂O₅ by weight relative to the total weight of the one or morevanadium oxides.

In one embodiment, the dehydrogenation catalyst has an average particlesize in the range of 20-160 μm.

In one embodiment, the dehydrogenation catalyst has an apparent particledensity in the range of 1-10 g/cm³.

In one embodiment, the dehydrogenation catalyst has a BET surface areain the range of 5-50 m²/g.

In one embodiment, the dehydrogenation catalyst is fluidizable and hasClass B powder properties in accordance with Geldart particleclassification.

According to a second aspect, the present disclosure relates to a methodfor producing the dehydrogenation catalyst of the present disclosure inany of its embodiments comprising i) mixing alumina with calcium oxideand a vanadyl coordination complex or salt in a solvent to form loadedcatalyst precursors, ii) reducing the loaded catalyst precursors with H₂gas to form reduced catalyst precursors, and iii) oxidizing the reducedcatalyst precursors with oxygen to form the dehydrogenation catalyst.

According to a third aspect, the present disclosure relates to a methoddehydrogenating an alkane to a corresponding olefin comprising flowingthe alkane through a reactor comprising a catalyst chamber loaded withthe dehydrogenation catalyst of the present disclosure in any of itsembodiments at a temperature in the range of 400-800° C. to form thecorresponding olefin and a reduced catalyst.

In one embodiment, the reactor is a fluidized bed reactor and thedehydrogenating is performed in a gas phase oxygen free environment.

In one embodiment, the alkane is propane and the corresponding olefin ispropylene.

In one embodiment, the method further comprises i) oxidizing at least aportion of the reduced catalyst in a gas phase oxygen environmentseparated from the catalyst chamber to regenerate the dehydrogenationcatalyst and ii) repeating the flowing and the oxidizing at least oncewith a less than 10% decrease in percent conversion of the alkane, aless than 10% decrease in selectivity for the olefin relative to a totalpercentage of products formed, or both.

In one embodiment, the dehydrogenation catalyst is present at an amountin the range of 0.05-1.0 g of catalyst per mL of alkane.

In one embodiment, the alkane is propane and the method has a propaneconversion of 10-80% at a reaction time of 5-60 seconds and atemperature of 500-700° C.

In one embodiment, the alkane is propane and the method has a propyleneselectivity of at least 60% relative to a total percentage of productsformed at a reaction time of 5-60 seconds and a temperature of 500-700°C.

In one embodiment, the alkane is propane and the method has a carbondioxide selectivity of no more than 40% relative to a total percentageof products formed at a reaction time of 5-60 seconds and a temperatureof 500-700° C.

In one embodiment, the alkane is propane, the dehydrogenation catalysthas a weight ratio of calcium oxide to alumina of 1:1, and the methodhas a propane conversion of at least 60% and a propylene selectivity ofat least 80% relative to a total percentage of products formed.

The foregoing paragraphs have been provided by way of generalintroduction, and are not intended to limit the scope of the followingclaims. The described embodiments, together with further advantages,will be best understood by reference to the following detaileddescription taken in conjunction with the accompanying drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

A more complete appreciation of the disclosure and many of the attendantadvantages thereof will be readily obtained as the same becomes betterunderstood by reference to the following detailed description whenconsidered in connection with the accompanying drawings, wherein:

FIG. 1A is the X-ray diffraction (XRD) patterns of the three preparedcatalyst samples VO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), andVO_(x)/CaO as well as their components V₂O₅, γ-Al₂O₃, and CaO.

FIG. 1B is the XRD patterns of the three prepared catalyst samplesVO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), and VO_(x)/CaO alonewith the peaks attributed to each of their components V₂O₅, γ-Al₂O₃, andCaO indicated.

FIG. 2A is the laser Raman spectroscopy spectra of the prepared catalystsamples VO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), andVO_(x)/CaO alone.

FIG. 2B is the laser Raman spectroscopy spectra of the prepared catalystsamples VO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), andVO_(x)/CaO) as well as their components V₂O₅, γ-Al₂O₃, and CaO.

FIG. 3A is the Fourier transform infrared (FTIR) absorption spectra ofthe prepared catalyst samples VO_(x)/CaO-γ-Al₂O₃ (1:4),VO_(x)/CaO-γ-Al₂O₃ (1:1), and VO_(x)/CaO as well as their componentsV₂O₅, γ-Al₂O₃, and CaO.

FIG. 3B is the FTIR absorption spectra of the prepared catalyst samplesVO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), and VO_(x)/CaOalone.

FIG. 4 is a scanning electron microscopy (SEM) image of the preparedVO_(x)/CaO-γ-Al₂O₃ (1:1) catalyst sample.

FIG. 5 is a scanning electron microscopy energy dispersive X-rayanalysis (SEM-EDX) image showing elemental mapping of vanadium theprepared VO_(x)/CaO-γ-Al₂O₃ (1:1) catalyst sample.

FIG. 6 is the temperature programmed reduction (TPR) profiles of theprepared VO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), andVO_(x)/CaO) catalyst samples.

FIG. 7 is the ammonia temperature programmed desorption (NH₃-TPD)profiles of the prepared VO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃(1:1), and VO_(x)/CaO) catalyst samples.

FIG. 8 is a comparison of the experimental data and fitted model ofammonia desorption during NH₃-TPD kinetics analysis for the differentprepared VO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), andVO_(x)/CaO) catalyst samples.

FIG. 9 is a graph of propane (C₃H₈) conversion and propylene (C₃H₆) andCO₂ selectivity and their error limits for the prepared catalysts inrepeated oxidative dehydrogenation of propane reaction runs at atemperature of 640° C., catalyst loading of 0.5 g, and a propaneinjection amount of 1.2 mL.

FIG. 10 is a graph of propane conversion for the prepared catalysts insuccessive propane injections without catalyst regeneration in oxidativedehydrogenation of propane reaction runs at a temperature of 640° C.,catalyst loading of 0.5 g, a propane injection amount of 1.2 mL, and acontact reaction time of 17 seconds.

FIG. 11 is a graph of propylene and CO₂ selectivity for the preparedcatalysts in successive propane injections without catalyst regenerationin oxidative dehydrogenation of propane reaction runs at a temperatureof 640° C., catalyst loading of 0.5 g, a propane injection amount of 1.2mL, and a contact reaction time of 17 seconds.

FIG. 12 is a graph of propane conversion for the prepared catalysts atdifferent temperatures in oxidative dehydrogenation of propane reactionruns at a catalyst loading of 0.5 g, a propane injection amount of 1.2mL, and a contact reaction time of 17 seconds.

FIG. 13 is a graph of propylene and CO₂ selectivity for the preparedcatalysts at different temperatures in oxidative dehydrogenation ofpropane reaction runs at a catalyst loading of 0.5 g, a propaneinjection amount of 1.2 mL, and a contact reaction time of 17 seconds.

FIG. 14 is a graph of propane conversion for the prepared catalysts atdifferent contact reaction times in oxidative dehydrogenation of propanereaction runs at a temperature of 640° C., catalyst loading of 0.5 g,and a propane injection amount of 1.2 mL.

FIG. 15 is a graph of propylene and CO₂ selectivity for the preparedcatalysts at different contact reaction times in oxidativedehydrogenation of propane reaction runs at a temperature of 640° C.,catalyst loading of 0.5 g, a propane injection amount of 1.2 mL.

FIG. 16 is a graph of propylene and CO₂ selectivity for the preparedcatalysts as a function of propane conversion for the prepared catalystsat constant temperature in oxidative dehydrogenation of propane reactionruns at a temperature of 640° C., catalyst loading of 0.5 g, and apropane injection amount of 1.2 mL.

DETAILED DESCRIPTION OF THE EMBODIMENTS

Referring now to the drawings. Embodiments of the present disclosurewill now be described more fully hereinafter with reference to theaccompanying drawings, in which some, but not all of the embodiments ofthe disclosure are shown.

According to a first aspect, the present disclosure relates to adehydrogenation catalyst, comprising i) a support material comprisingalumina modified by calcium oxide, and ii) a catalytic materialcomprising one or more vanadium oxides disposed on the support material,wherein the dehydrogenation catalyst comprises 5-20% of the one or morevanadium oxides by weight relative to the total weight of thedehydrogenation catalyst.

Vanadium oxide (e.g., V₂O₅—vanadia) is considered to be one of the mostimportant and useful metals to be used as a catalyst due to its physicaland chemical properties, and catalysis is the most dominantnon-metallurgical use of vanadia. The catalytic activity of vanadia isattributed to its reducible nature and its ability to easily change itsoxidation state from V⁺³ to V⁺⁵. It is generally accepted that V⁺⁵ isthe highly active initial state of the catalyst in a cycle of oxidativedehydrogenation. Vanadium oxide catalysts have been used in manyindustrial and lab scale catalytic reactions and processes. In manycases, vanadia catalysts are doped with promoters to improve theiractivity or selectivity, while various supports are used to improvemechanical strength, thermal stability, longevity, and/or catalyticperformance.

As used herein, a catalyst support material refers to material, usuallya solid with a high surface area, to which a catalyst is affixed. Thereactivity of heterogeneous catalyst and nanomaterial based catalystsoccurs at the surface atoms. Thus, great effort is made herein tomaximize the surface of a catalyst by distributing it over the support.The support may be inert or participate in the catalytic reactions. Thesupport materials used in catalyst preparation play a role indetermining the physical characteristics and performance of thecatalysts. Typical supports include various kinds of carbon, alumina andsilica. In a preferred embodiment, the dehydrogenation catalyst of thepresent disclosure comprises an alumina support material, preferably acalcium oxide modified alumina support material.

As used herein, alumina refers to aluminum oxide, a chemical compound ofaluminum and oxygen with the chemical formula Al₂O₃. Aluminum oxide iscommonly called alumina and may also be referred to as aloxide, aloxite,or alundum. It is the most commonly occurring of several aluminum oxidesand specifically identified as aluminum (III) oxide. It commonly occursin its crystalline polymorphic phase α-Al₂O₃ which composes the mineralcorundum, the most thermodynamically stable form of aluminum oxide.Al₂O₃ is significant in its use to produce aluminum metals and noted forits high melting point. In one embodiment, the catalytic material isloaded on an inert alumina support. Exemplary inert alumina based inertmaterials include, but are not limited to aluminum oxide, alumina,alumina monohydrate, alumina trihydrate, alumina silica, bauxite,calcined aluminum hydroxides such as gibbsite, bayerite and boehmite aswell as calcined hydrotalcite and the like.

In one embodiment, the alumina support material may be comprised of aplurality of different crystallographic phases. In the most common andthermodynamically stable form, corundum, the oxygen ions nearly form ahexagonal close-packed structure with aluminum ions filling two-thirdsof the octahedral interstices. Each Al³⁺ center is octahedral. In termof its crystallography, corundum adopts a trigonal Bravais lattice andits primitive cell contains two formula units of aluminum oxide.Aluminum oxide also exists in other phases, including the transitioncubic γ and η phases, the monoclinic θ phase, the hexagonal χ phase, theorthorhombic κ phase and the transition δ phase that can be tetragonalor orthorhombic. Each has unique crystal structure and properties. Inthe present disclosure, aluminum oxide or alumina may refer to Al₂O₃having an a polymorph, a γ polymorph, a η polymorph, a θ polymorph, a χpolymorph, a κ polymorph and a δ polymorph or mixtures thereof,preferably a γ polymorph. In a preferred embodiment, the alumina of thepresent disclosure consists substantially of γ-Al₂O₃, preferably greaterthan 75% by weight relative to the total weight of alumina, preferablygreater than 80%, preferably greater than 85%, preferably greater than90%, preferably greater than 95%, preferably greater than 98%,preferably greater than 99% by weight relative to the total weight ofthe alumina. In at least one embodiment, the alumina support materialconsists essentially of γ-alumina (γ-Al₂O₃).

Alumina, especially γ-Al₂O₃ is used for its very high surface area onwhich active metal atoms/crystallites can spread out as reactive sites,but also for its enhancement of productivity and/or selectivity throughmetal-support interaction and spillover/reverse-spillover phenomena. Inreactions, γ-Al₂O₃ must retain as much high surface area duringreaction. Additives and/or modifiers and additional supports markedlyincrease the thermal stability of the support, effect acidity and activesite metal support-interactions and prevent the loss of surface areaunder thermal reaction conditions.

In a preferred embodiment, the support material comprising alumina ismodified by calcium oxide. As used herein, calcium oxide (CaO) alsoknown as quicklime or burnt lime refers to a widely used chemicalcompound. It is a white, caustic, alkaline, crystalline solid at roomtemperature. The broadly used term “lime” connotes calcium containinginorganic materials, in which carbonates, oxides and hydroxides ofcalcium, silicon, magnesium, aluminum, and iron predominate. Bycontrast, “quicklime” specifically applies to the single chemicalcompound calcium oxide. Calcium oxide which survives processing withoutreacting is often termed “free lime”. Quicklime is relativelyinexpensive. Both it and a chemical derivative (calcium hydroxide, ofwhich quicklime is the base anhydride) are important commoditychemicals.

In a preferred embodiment, the dehydrogenation catalyst of the presentdisclosure comprises 10-85% of alumina by weight relative to the totalweight of the dehydrogenation catalyst, preferably 15-80%, preferably25-75%, preferably 35-74%, preferably 45-72% of alumina by weightrelative to the total weight of the dehydrogenation catalyst. In apreferred embodiment, the dehydrogenation catalyst of the presentdisclosure comprises 5-80% of calcium oxide by weight relative to thetotal weight of the dehydrogenation catalyst, preferably 10-75%,preferably 15-60%, preferably 18-46%, preferably 20-40% of calcium oxideby weight relative to the total weight of the dehydrogenation catalyst.In a preferred embodiment, the dehydrogenation catalyst of the presentdisclosure has a weight ratio of calcium oxide to alumina in the rangeof 1:0.2 to 1:8, preferably 1:0.4 to 1:7.5, preferably 1:0.5 to 1:7,preferably 1:0.6 to 1:6.5, preferably 1:0.8 to 1:6, preferably 1:0.9 to1:5, preferably 1:1 to 1:4, preferably 1:1.5 to 1:3, preferably 1:1.75to 1:2.5. In a more preferred embodiment, the dehydrogenation catalystof the present disclosure has a weight ratio of calcium oxide to aluminain the range of 1:0.5 to 1:4, preferably 1:0.75 to 1:2, most preferably1:1. In a most preferred embodiment, the dehydrogenation catalyst of thepresent disclosure has a weight ratio of calcium oxide to alumina of1:1.

It is equally envisaged that the dehydrogenation catalyst of the presentdisclosure may be adapted to incorporate additional support materialsand additional additives such as phase transformation stabilizers. Insome embodiments, these additional support materials and additionaladditives may be used in addition to, or in lieu of alumina and/orcalcium oxide. Exemplary additional support materials include, but arenot limited to oxides such as, SiO₂, TiO₂, ZrO₂, CeO, NbO₅, MgO andzeolites. Exemplary additional thermal stabilizer additives include, butare not limited to, the elements La, Ce, Ba, Sr, Sm, Si, Pr and P. Whenlanthanum is used as an additive, the formation of lanthanum aluminatecan decrease the surface energies of γ-Al₂O₃ lowering the driving forcefor sintering and stabilizing bulk phase transformation. In certainembodiments, the dehydrogenation catalyst of the present disclosurecomprises less than 5% of additional additives, such as elementallanthanum, by weight relative to the total weight of the dehydrogenationcatalyst, preferably 0.1-3.0% of additional additives by weight relativeto the total weight of the dehydrogenation catalyst, preferably0.5-2.0%, preferably 0.75-1.5%, preferably 0.8-1.1%, or about 1.0% ofadditional additives by weight relative to the total weight of thedehydrogenation catalyst.

In a preferred embodiment, the dehydrogenation catalyst of the presentdisclosure comprises a catalytic material disposed on the supportmaterial, wherein the catalytic material comprises one or more vanadiumoxides. As used herein, “disposed on” or “impregnated” describes beingcompletely or partially filled throughout, saturated, permeated and/orinfused. The catalytic material may be affixed on one or more surfacesof the support material the catalytic material may be affixed on anouter surface of the support material or within pore spaces of thesupport material. The catalytic material may be affixed to the supportmaterial in any reasonable manner, such as physisorption orchemisorption and mixtures thereof. In one embodiment, greater than 10%of the surface area (i.e. surface and pore spaces) of the supportmaterial is covered by the catalytic material, preferably greater than15%, preferably greater than 20%, preferably greater than 25%,preferably greater than 30%, preferably greater than 35%, preferablygreater than 40%, preferably greater than 45%, preferably greater than50%, preferably greater than 55%, preferably greater than 60%,preferably greater than 65%, preferably greater than 70%, preferablygreater than 75%, preferably greater than 80%, preferably greater than85%, preferably greater than 90%, preferably greater than 95%,preferably greater than 96%, preferably greater than 97%, preferablygreater than 98%, preferably greater than 99%. In preferred embodiments,the vanadium or vanadium oxide comprising catalytic material ishomogeneously distributed or dispersed throughout the support materialand on the surface of the support material. In preferred embodiments,this quality of the dispersion can be verified scanning electronmicroscopy (SEM) and/or energy dispersive X-ray analysis (EDX) providingelemental mapping, preferably vanadium elemental mapping. In otherembodiments the catalytic material may form localized clusters amongstthe support material, form oxide species with the support catalyst orform layers of the catalytic material and vanadium species amongst thesupport material, or be heterogeneously disposed on the support materialand its surfaces and mixtures thereof.

In a preferred embodiment, the catalytic material comprises one or morevanadium oxides. In terms of the present disclosure, vanadium oxide mayrefer to vanadium (II) oxide (vanadium monoxide, VO), vanadium (III)oxide (vanadium sesquioxide or trioxide, V₂O₃), vanadium (IV) oxide(vanadium dioxide, VO₂), vanadium (V) oxide (vanadium pentoxide, V₂O₅).Vanadium oxide may also refer to a vanadate, a compound containing onoxoanion of vanadium generally in its highest oxidation state of ⁺5. Thesimplest vanadate ion is the tetrahedral orthovanadate VO₄ ³⁻ anion.Exemplary vanadate ions include, but are not limited to, VO₄ ³⁻, V₂O₇⁴⁻, V₃O₉ ³⁻, V₄O₁₂ ⁴⁻, V₅O₁₄ ³⁻ and the like. In addition to theseprincipal oxides of vanadium, various other distinct phases exist.Phases with the general formula V_(n)O_(2n+1), wherein n is a wholenumber greater than zero exist between V₂O₅ (vanadium (V) species) andvanadium (IV) species. Examples of these phases include V₃O₇, V₄O₉ andV₆O₁₃. Phases with the general formula V_(n)O_(2n−1), wherein n is awhole number greater than zero exist between vanadium (IV) species andV₂O₃ (vanadium (III) species). Termed Magneli phases, they are examplesof crystallographic shear compounds based on rutile structure. Examplesof Magneli phases include V₄O₇, V₅O₉, V₆O₁₁, V₇O₁₃ and V₈O₁₅. Manyvanadium oxygen phases are non-stoichiometric. In a preferredembodiment, the dehydrogenation catalyst of the present disclosurecomprises 5-20% of the one or more vanadium oxides by weight relative tothe total weight of the dehydrogenation catalyst, preferably 6-18%,preferably 7-16%, preferably 8-14%, preferably 9-12%, or about 10% ofthe one or more vanadium oxides by weight relative to the total weightof the dehydrogenation catalyst.

In a preferred embodiment, the one or more vanadium oxides are of theformula V_(x)O_(y) wherein x=1-4, preferably 1-3, more preferably 1-2and y=2-10, preferably 2-5. In a preferred embodiment, the one or morevanadium oxides are at least one selected from the group consisting ofV₂O₅, VO₂ and V₂O₃. V₂O₅ or vanadium (V) oxide or vanadium pentoxide isan inorganic compound that due to its high oxidation state is both anamphoteric oxide and an oxidizing agent. V₂O₅ is characterized by itsvaluable redox properties as V₂O₅ is easily reduced to the stablevanadium (IV) species. In certain embodiments, the dehydrogenationcatalyst comprises at least 50% of V₂O₅ by weight relative to the totalweight of the one or more vanadium oxides, preferably greater than 60%,preferably greater than 70%, preferably greater than 80%, preferablygreater than 85%, preferably greater than 90%, preferably greater than95%, preferably greater than 96%, preferably greater than 97%,preferably greater than 98%, preferably greater than 99% of V₂O₅ byweight relative to the total weight of the one or more vanadium oxides,such as, for example 50-90% by weight V₂O₅, preferably 75-80% V₂O₅, morepreferably 85-90% V₂O₅, even more preferably at least 90-95% V₂O₅, mostpreferably 95-99.9% V₂O₅ relative to the total weight of the one or morevanadium oxides. In certain embodiments, the dehydrogenation catalyst ofthe present disclosure consists essentially of V₂O₅ and is substantiallyfree of V₂O₃ and VO₂. In some embodiments, the dehydrogenation catalystof the present disclosure is substantially free of V₂O₃ and comprises amixture of at least 50% V₂O₅ by weight relative to the total weight ofthe one or more vanadium oxides, with the balance substantiallycomprising VO₂.

The different vanadia phases that can be present in supported vanadiaoxide catalysts as well as the distribution among the various vanadiumoxide structures can depend on the synthesis method, the vanadiumprecursor, solvent, calcination temperature, vanadium oxide loading,oxide support, etc. At loadings below “monolayer coverage” isolated andoligomerized surface VO₄ species may be present on the oxide support.The surface VO₄ species may possess up to three different oxygen atomsincluding, but not limited to, oxygen atoms forming a vanadyl group(V═O), oxygen atoms bridging two vanadia atoms (V—O—V), and oxygen atomsbridging a vanadia atom and oxide support cation (V—O-support).Depending on the vanadia surface density as well as the supportmaterial, a vanadia “monolayer coverage” may be reached. A “monolayer”refers to a single, closely packed layer of atoms or molecules, here theone or more vanadium oxides. As used herein, “monolayer coverage” refersto the completion of a 2D surface of vanadium oxide overlayer on thealumina support, and the surface becomes saturated before 3D vanadiumoxide and/or V₂O₅ crystallites start to form and grow subsequently. In apreferred embodiment, the vanadium loading is below the monolayercoverage and the VO_(x) species in the catalytic material are highlydispersed forming an amorphous phase on the γ-Al₂O₃ and CaO supportsurface. Alternatively, the monolayer coverage may be thought of as theminimum amount of single vanadium and/or vanadium oxide atoms ormolecules to cover exactly 100% of the surface area (surface and porespaces) of the support material uniformly. In a preferred embodiment,the monolayer coverage of the dehydrogenation catalyst of the presentdisclosure corresponds to 5-20 vanadium atoms per nm² of support,preferably 6-15 atoms/nm², preferably 7-10 atoms/nm², preferably 8-9vanadium atoms per nm² of support. In certain embodiments, V₂O₅crystallites may be present at vanadium oxide loadings below monolayercoverage when a precursor vanadium salt is not well dispersed over thesupport during synthesis or when a weak interaction exists between thevanadium oxide and the support. In one embodiment, the one or morevanadium oxides may form a crystalline phase on the surface of thelanthanum modified alumina support material, preferably a V₂O₅crystalline phase. At high enough loadings, greater than monolayercoverage, vanadium oxide nanocrystals or nanoparticles having an averageparticle size of 1-100 nm, preferably 4-80 nm, preferably 10-60 nm,preferably 20-40 nm may be present on the surface of the catalystsupport. In certain embodiments, the different surface vanadia speciesmay be identified by techniques including, but not limited to, Ramanspectroscopy, Fourier transform infrared spectroscopy (FT-IR), UV-visspectroscopy, X-ray powder diffraction (XRD) and the like. In apreferred embodiment, the one or more vanadium oxides form an amorphousphase on the surface of the support material. Alternatively, it isenvisaged that the catalytic material comprising one or more vanadiumoxides of the present disclosure forms a crystalline phase on thesupport surface. In other embodiments, the catalytic material maydisplay a mixed amorphous and crystalline phase.

In certain embodiments, the catalytic material comprises one or morevanadium oxides and may optionally further comprise a promoter. As usedherein, a promoter refers to an additive to improve catalystperformance. Metal promoters such as for example niobium may function toisolate active species (i.e. VO_(x), more preferably V₂O₅) and to formsecondary metallic oxides (i.e. Nb₂O₅) on support surface. Furthermore,the addition of promoters to the catalytic material blocks acid siteswhich decreases the total acidity of the dehydrogenation catalyst. Incertain embodiments, the decrease in acidity and increase in basicitymay facilitate desorption of substrates from the dehydrogenationcatalyst surface, preventing further oxidation, such as, for example theundesirable combustion to carbon oxides (CO_(x)) in the oxidativedehydrogenation of light alkanes such as ethane and propane. In apreferred embodiment, the dehydrogenation catalyst of the presentdisclosure may further comprise 1.0-5.0% of promoter by weight relativeto the total weight of the dehydrogenation catalyst, preferably1.5-4.0%, preferably 2.0-3.75%, preferably 3.0-3.5%, or about 3.25% ofpromoter by weight relative to the total weight of the dehydrogenationcatalyst. Exemplary promoters include, but are not limited to, metallicpromoters (Nb, Cr, Mo, Ta, W), alkali promoters (Li, K, Rb) and halidepromoters (Cl) and mixtures thereof. In preferred embodiments, thevanadium or vanadium oxide and promoter or promoters are homogeneouslydistributed throughout the catalyst support. In other embodiments thepromoter may form localized clusters amongst the vanadium, form promoteroxide species with the support catalyst, form layers of promoter andvanadium species, or be disposed on the vanadium oxide species andmixtures thereof.

In a preferred embodiment, the present disclosure provides fluidizabledehydrogenation catalysts for oxidative dehydrogenation (ODH) of alkanespreferably in reactors having a fluidized bed design. As used herein“fluidizable” refers to the ability to undergo fluidization which refersto a process similar to liquefaction whereby a granular material isconverted from a static solid-like to a dynamic fluid-like state. Theprocess occurs when a fluid (liquid or gas) is passed up through thegranular material. A fluidized bed is formed when a quantity of a solidparticulate substance is placed under appropriate conditions to cause asolid/fluid mixture to behave as a fluid. This is usually achieved bythe introduction of pressurized fluid through the particulate medium.This results in the medium then having many properties andcharacteristics of normal fluids, such as the ability to free flow undergravity, or to be pumped using fluid type technologies. Fluidized bedtypes can be broadly classified by their flow behavior including, butnot limited to, stationary or bubbling fluidized beds, circulatingfluidized beds (CFB), vibratory fluidized beds, transport or flashreactor (FR), and annular fluidized beds (AFB).

In certain fluidized bed reactors, the catalyst pellets lie on a grateat the bottom of the reactor. Reactants are continuously pumped into thereactor through a distributor causing the bed to become fluidized.During the fluidization, the catalyst pellets are converted from astatic solid like state to a dynamic fluid like state. The bed'sbehavior after initial fluidization depends on the state of thereactant. If it is a liquid the bed expands uniformly with an increasedupward flow of the reactant, resulting in a homogeneous fluidization. Ifthe reactant is a gas, the bed will be non-uniform because the gas formsbubbles in the bed, resulting in aggregative fluidization. In terms ofthe present disclosure, the fluidization may be homogeneous oraggregative. In certain embodiments, the reactant or feed is preferablya light alkane including, but not limited to, ethane, propane and butane(including n-butane and isobutene), all of which are gases and hence, anaggregative fluidization may be probable.

Properties and parameters for determining the fluidizability,reducibility, and oxygen carrying capacity of a catalyst can be bothmeasured and calculated. The average particle size and the particle sizedistribution can be measured, for example, using a Mastersizer 2000 fromMalvern Instruments. For spherical or substantially sphericaldehydrogenation catalyst particles, average particle size refers to thelongest linear diameter of the dehydrogenation catalyst particles. In apreferred embodiment, the dehydrogenation catalyst of the presentdisclosure in any of its embodiments has an average particle size in therange of 20-160 μm, preferably 30-150 μm, preferably 40-120 μm,preferably 50-100 μm, more preferably 60-80 μm. In one embodiment, theparticle size distribution of the dehydrogenation catalyst of thepresent disclosure is 10-200 μm and greater than 75% of the particleshave a particle size of 40-120 μm, preferably greater than 80%,preferably greater than 85%, more preferably greater than 90% have aparticle size of 40-120 μm. In another embodiment, the dehydrogenationcatalyst of the present disclosure has a particle size distributionranging from 33% of the average particle size to 133% of the averageparticle size, preferably 50-130%, preferably 60-125%, preferably80-100%, preferably 90-110%, preferably 95-105% of the average particlesize. In one embodiment, the dehydrogenation catalyst particles of thepresent disclosure are monodisperse, having a coefficient of variationor relative standard deviation, expressed as a percentage and defines asthe ratio of the particle size standard deviation (σ) to the particlemean size (μ) multiplied by 100 of less than 25%, preferably less than20%, preferably less than 15%, preferably less than 12%, preferably lessthan 10%, preferably less than 8%, preferably less than 6%, preferablyless than 5%.

As used herein, the apparent particle density refers to the mass of thecatalyst divided by the volume that it occupies. The apparent particledensity can be assessed using a CREC-established method. In the method,a known amount of catalyst is introduced to a flask. The flask is filledwith isopropanol and the apparent particle density (AD) is calculatedusing the following equation formula (I).

$\begin{matrix}{{AD} = \frac{W_{cat}}{V_{T} - V_{isopropanol}}} & (I)\end{matrix}$Where AD is the apparent particle density (g/cm³), W_(cat) is thecatalyst weight, V_(T) is the flask volume and V_(isopropanol) is thevolume of isopropanol calculated as the ratio of the weight ofisopropanol needed to fill the flask and the density of isopropanol. Ina preferred embodiment, the dehydrogenation catalyst of the presentdisclosure in any of its embodiments has an apparent particle density of1.0-10.0 g/cm³, 1.1-5.0 g/cm³, preferably 1.25-4.0 g/cm³, preferably1.5-3.5 g/cm³, more preferably 1.8-3.2 g/cm³.

In some embodiments, with the calculated average particle size andparticle apparent density values, the fluidization regime of thedehydrogenation catalyst particles of the present disclosure can bedetermined using Geldart's powder classification chart. Geldart groupspowders into four “Geldart Groups” or “Geldart Classes”. The groups aredefined by solid-fluid density difference and particle size. Designmethods for fluidized beds can be tailored based upon a particle'sGeldart Group. For Geldart Group A the particle size is between 20 and100 μm and the particle density is typically less than 1.4 g/cm³. Priorto the initiation of a bubbling bed phase, beds from these particleswill expand by a factor of 2 to 3 at incipient fluidization, due to tadecreased bulk density. Most powder-catalyzed beds utilize this group.For Geldart Group B the particle size lies between 40 and 500 μm and theparticle density is between 1.4-4 g/cm³. Bubbling typically formsdirectly at incipient fluidization. For Geldart Group C the groupcontains extremely fine and consequently the most cohesive particles.With a particle size of 20 to 30 μm, these particles fluidize under verydifficult to achieve conditions, and may require the application of anexternal force, such as mechanical agitation. For Geldart Group D theparticles in this regime are above 600 μm and typically have highparticle densities. Fluidization of this group requires very high fluidenergies and is typically associated with high levels of abrasion.Additionally, these particles are usually processed in shallow beds orin the spouting mode. The dehydrogenation catalyst of the presentdisclosure is preferably fluidizable and may be classified as a GeldartGroup A powder, a Geldart Group B powder, a Geldart Group C powder or aGeldart Group D powder, preferably as a Geldart Group B powder. In atleast one preferred embodiment, the dehydrogenation catalyst particlesdisplay a Geldart Group B powder property, which is highly fluidizableunder ODH conditions. Large particles, such as those under Geldart GroupD, may limit the gas phase reactant access to the inner layers of thecatalyst. As a result, using smaller particles can minimize thediffusional resistance and reduction/oxidation rates can be maximized.On the other hand, very small particles, such as those under Geldart'sGroup C, can cause fluidization problems, channeling and loss of fines.

The Brunauer-Emmet-Teller (BET) theory aims to explain the physicaladsorption of gas molecules on a solid surface and serves as the basisfor an important analysis technique for the measurement of the specificsurface area of a material. Specific surface area is a property ofsolids which is the total surface area of a material per unit of mass,solid or bulk volume, or cross sectional area. In a preferredembodiment, the dehydrogenation catalyst of the present disclosure inany of its embodiments has a BET surface area in the range of 5-50 m²/g,preferably 10-45 m²/g, preferably 11-40 m²/g, preferably 12-30m²/g/preferably 13-28 m²/g, preferably 14-25 m²/g.

The catalytic activity of many oxides in various processes is due totheir Lewis and Bronsted acidities. In addition to effects on surfacearea, catalyst modifications (i.e. the modification of alumina with CaO)may also decrease the surface acidity and metal-support interactions ofthe catalyst, thereby enhancing olefin selectivity in oxidativedehydrogenation reactions and reducing coke (CO_(x)) formation. Thecatalyst acidity plays a role in metal support interactions that affectVO_(x) reducibility. The reducibility may impact catalyst activity andselectivity by providing O₂ for oxidation and high acidity not favoringselective oxidation. A number of techniques have been developed for thecharacterization of acid-base surface properties of catalysts. Theadsorption of volatile amines including, but not limited to, ammonia(NH₃), pyridine (C₅H₅N), n-butylamine (CH₃CH₂CH₂CH₂NH₂), quinolone(C₉H₇N) and the like is often used to determine the acid siteconcentration of solid catalysts. The amount of the base remaining onthe surface after evacuation is considered chemisorbed and serves as ameasure of the acid site concentration. The adsorbed base concentrationas a function of evacuation temperature can give a site strengthdistribution. Another means of determining the site strengthdistribution is calorimetry or the temperature-programmed desorption(TPD).

Ammonia or NH₃-TPD experiments are used to determine the total acidityof the catalyst. TPD can further give an idea about metal-supportinteractions by modeling NH₃ desorption kinetics and be used todetermine the strength of acid sites available on the catalyst surface.In a preferred embodiment, the dehydrogenation catalyst of the presentdisclosure in any of its embodiments has a total acidity in the range of0.1-5.0 mmol of NH₃ per gram of catalyst, preferably 0.2-4.0 mmol of NH₃per gram of catalyst, preferably 0.3-3.0 mmol of NH₃ per gram ofcatalyst, preferably 0.4-2.8 mmol of NH₃ per gram of catalyst,preferably 0.5-2.5 mmol of NH₃ per gram of catalyst when measured with aheating rate of 5-20° C./min, preferably 10-15° C./min. In a preferredembodiment, the dehydrogenation catalyst of the present disclosure has alower acidity than pure alumina. In a preferred embodiment, thedehydrogenation catalyst of the present disclosure has an energy of NH₃desorption established by NH₃-TPD kinetic analysis and an indicator ofactive site metal-support interactions in the range of 10-100 kJ/mol,preferably 25-75 kJ/mol, preferably 30-60 kJ/mol, preferably 35-50kJ/mol. In one embodiment, the inclusion of greater amounts of CaO tothe support material comprising alumina modified by calcium oxide maydecrease the total acidity of the dehydrogenation catalyst of thepresent disclosure, increase the oxygen carrying capacity of thedehydrogenation catalyst of the present disclosure and decrease theactivation energy of ammonia desorption of the dehydrogenation catalystof the present disclosure relative to a substantially similar catalystlacking CaO modified support material or comprising a lower weight ratioof calcium oxide to alumina. In addition, the effects of intermediatecatalyst acidity, moderate active site metal-support interactions andmoderate active site metal-support interactions may favor productselectivity in oxidative dehydrogenation reactions.

According to a second aspect, the present disclosure relates to a methodfor producing the dehydrogenation catalyst of the present disclosure inany of its embodiments, comprising i) mixing alumina with calcium oxideand a vanadyl coordination complex or salt in a solvent to form loadedcatalyst precursors, ii) reducing the loaded catalyst precursors with H₂gas to form reduced catalyst precursors, and iii) oxidizing the reducedcatalyst precursors with oxygen to form the dehydrogenation catalyst.

Two main methods are typically used to prepare supported catalysts. Inthe impregnation method, the solid support or a suspension of the solidsupport is treated with a solution of a precatalyst (for instance ametal salt or metal coordination complex), and the resulting materialthen activated under conditions that will convert the precatalyst to amore active state, such as the metal itself or metal oxides of themetal. In such cases, the catalyst support is usually in the form ofpellets or spheres. Alternatively, supported catalysts can be preparedfrom homogenous solution by co-precipitation. In terms of the presentdisclosure, it is envisaged that the dehydrogenation catalyst may beformed by an impregnation method or a co-precipitation method,preferably by an impregnation method, preferably by an impregnationmethod through soaking with an excess solvent. Supports are usuallythermally very stable and withstand processes required to activateprecatalysts. For example, many precatalysts are activated by exposureto a stream of hydrogen or air (oxygen) at high temperatures,additionally many precatalysts may be activated and/or reactivated byoxidation-reduction cycles, again at high temperatures.

In one step of the process, alumina is mixed with calcium oxide and avanadyl coordination complex or salt in a solvent to form loadedcatalyst precursors. In one embodiment, the unmodified alumina andcalcium oxide supports before metal loading may be optionally initiallydried and/or calcined to remove moisture and other volatile compounds.In one embodiment, the preemptive calcining may be performed at atemperature of 300-600° C., preferably 400-550° C., or about 550° C. fora period of up to 8 hours, preferably up to 6 hours, preferably up to 4hours, or about 4 hour. The precalcining may be performed at atemperature of 600-800° C., preferably 650-750° C., or about 725° C. fora period of up to 8 hours, preferably up to 6 hours, or about 4 hours.

The manner in which the vanadium oxide is deposited onto a support canhave an influence on the properties of the active component in the finalcatalyst. Typically the main method of dispersing vanadium oxide onsupport materials is the classic incipient wetness impregnation methodin a solvent where the vanadium salt is soluble. The impregnation methodis performed by contacting the support with a certain volume of solutioncontaining the dissolved vanadium oxide precursor. If the volume of thesolution is either equal to or less than the pore volume of the support,the technique is referred to as incipient wetness. This particularsynthesis route can show a broad variation of vanadium oxide surfacespecies at all loadings, particularly loadings below monolayer coverage,depending on the synthesis conditions. In one embodiment, this methodmay lead to the formation of three-dimensional V₂O₅ nanoparticles, evenat low vanadium oxide loadings. In another embodiment, this method maylead to the formation of an amorphous vanadium oxide phase on thesurface of the support.

In a preferred embodiment, the loaded catalyst precursors are preparedby an incipient wetness method of impregnation. The alumina support canbe immersed in a solution comprising calcium oxide and vanadium and/or avanadium salt or coordination complex. In one embodiment, the vanadiumsalt or coordination complex may be a vanadium (IV), vanadium (V) orvanadium (III) salt. Exemplary vanadium salts or coordination complexesinclude, but are not limited to, ammonium metavanadate in mixtures ofwater and oxalic acid or methanol and oxalic acid, vanadium (III)acetylacetonate (V(AcAc)₃) or vanadyl acetylacetonate (VO(AcAc)₂) intoluene, VO(iPrO)₃, VO(OC₂H₅)₃, or VO(OC₂H₇)₃ in 2-propanol, as well asvanadyl sulfate, vanadium pentoxide, vanadium (III) chloride, vanadiumoxytripropoxide, tetrakis(diethylamido)vanadium (IV), vanadium (IV)chloride, vanadium (III) chloride tetrahydrofuran complex, vanadium (V)oxychloride, vanadium (V) oxyfluoride, and the like. Preferably, thevanadium salt or coordination complex is vanadium (III) acetylacetonate(V(AcAc)₃) or vanadyl acetylacetonate (VO(AcAc)₂), most preferablyvanadyl acetylacetonate (VO(AcAc)₂). The vanadium salt is preferablyphosphorous free. In a preferred embodiment, the solvent is a polarprotic solvent. Exemplary polar protic solvents include, but are notlimited to, formic acid, n-butanol, isopropanol, n-propanol, ethanol,methanol, acetic acid, water and mixtures thereof, preferably thesolvent is ethanol. It is equally envisaged that the present method maybe adapted to incorporate non-polar solvents including, but not limitedto, pentane, cyclopentane, hexane, cyclohexane, benzene, toluene,1,4-dioxane, chloroform, diethyl ether and dichloromethane, as wellpolar aprotic solvents including, but not limited to, tetrahydrofuran,ethyl acetate, acetone, dimethylformamide, acetonitrile, dimethylsulfoxide, nitromethane, propylene carbonate and mixtures thereof.

In a preferred embodiment the vanadium salt is vanadyl acetylacetonateVO(AcAc)₂ and the solvent is ethanol. In a preferred embodiment thesolution has a vanadium concentration of 0.01-1.0 M, preferably 0.05-0.5M, preferably 0.1-0.25 M, preferably 0.125-0.2 M, or about 0.15 M. In apreferred embodiment, the mixing of the alumina support material withthe calcium oxide and vanadyl coordination complex or salt in a solventis performed at a temperature of 20-40° C., preferably 20-30° C., orabout 25° C. for a period of less than 48 hours, preferably less than 36hours, preferably less than 24 hours, preferably less than 18 hours,preferably less than 12 hours, preferably less than 10 hours andoptionally with stirring and/or ultrasonication for 1-60 minutes,preferably 5-30 minutes, preferably 10-20 minutes to achieve ahomogeneous mixture. After mixing the solution can be filtered andseparated from the solvent to provided loaded catalyst precursors.

In another embodiment, it is equally envisaged that the method may beadapted to other means of dispersing and depositing the vanadium oxideon the support material. Both adsorption from solution (i.e. grafting)based on attaching vanadia from the solution through reaction withhydroxyl groups on the surface of the support and ion exchange methodspermitting the ionic vanadium oxide species present in an aqueoussolution to be electrostatically attracted by charged sites of thesupport surface have been used. Exemplary other means include, but arenot limited to, vapor-fed flame synthesis, flame spray pyrolysis,sputter deposition, atomic layer deposition and chemical vapordeposition (CVD). For example, chemical vapor deposition (CVD) usesvolatile molecular metal precursors (i.e. O═VCl₃, O═V(OC₂H₅)₃ orO↑V(OiPr)₃) to modify oxide support surface and provide a way to controlthe dispersion of the active sites.

In certain embodiments, in addition to the methods employed to dispersevanadium oxide material on different supports, the drying and/orcalcination used for the fixation of the vanadia may be a crucial stepof the catalyst preparation due to the conversion of the initialvanadium species that may result in a broad variety of V_(x)O_(y)species from a nominally simple impregnation process. At highcalcination temperatures, mixed oxide compounds or solid solutions canbe formed with some oxide supports (i.e. AlVO₄). In a preferredembodiment, the loaded catalyst supports are dried before the reductionand the oxidation at room temperature and following natural dryingbefore the reduction and the oxidation at a temperature of up to 300°C., preferably up to 250° C., preferably up to 200° C., preferably up to175° C., preferably up to 150° C., preferably up to 125° C., preferablyup to 100° C. for a period of up to 60 hours, preferably up to 48 hours,preferably up to 36 hours, preferably up to 24 hours, preferably up to12 hours.

In one step of the process the loaded catalyst precursors are reducedwith H₂ gas to form reduced catalyst precursors. As used herein,reduction refers to the gain of electrons or a decrease in oxidationstate by a molecule, atom or ion. In a preferred embodiment, the loadedcatalyst precursors are reduced under a flow of hydrogen gas comprising1-40% H₂, preferably 2-20% H₂, preferably 4-18% H₂, preferably 6-16% H₂,preferably 8-14% H₂, or about 10% H₂ as a molar percentage and 60-99%inert gas, preferably 70-95% inert gas, preferably 80-94% inert gas,preferably 85-92% inert gas, or about 90% inert gas as a molarpercentage. Exemplary inert gases include nitrogen (N₂) and argon (Ar),preferably argon. In a preferred embodiment, the reduction underhydrogen gas flow is performed at a temperature of 300-800° C.,preferably 350-750° C., preferably 400-700° C., preferably 425-650° C.,preferably 450-600° C., preferably 475-550° C., or about 500° C. for aperiod of 1-18 hours, preferably 2-12 hours, preferably 4-8 hours, orabout 6 hours. In certain embodiments, the reduction of the loadedcatalyst precursors may be performed in a fluidized bed reactor.

In one step of the process the reduced catalyst precursors are oxidizedwith oxygen to form the dehydrogenation catalyst of the presentdisclosure in any of its embodiments. As used herein, oxidation refersto the loss of electrons or an increase in oxidation state by amolecule, atom or ion. Oxidation reactions are commonly associated withthe formation of oxides from oxygen molecules. Oxygen itself is the mostversatile oxidizer. In a preferred embodiment, the reduced catalystprecursors are oxidized under air flow comprising 20-25% O₂, preferably20.5-22% O₂, or about 21% O₂ as a molar percentage and 75-80% N₂,preferably 77-79% N₂, or about 78% N₂ as a molar percentage. In apreferred embodiment, the oxidation under air flow or calcination underair flow is performed at a temperature of 300-700° C., preferably350-650° C., preferably 400-600° C., preferably 425-575° C., preferably450-550° C., preferably 475-525° C., or about 500° C. for a period oftime of 1-12 hours, preferably 2-8 hours, preferably 3-6 hours, or about4-5 hours. In certain embodiments, obtaining the oxide catalyst formwill be accompanied by a characteristic yellow color or color changeindicating the presence of V₂O₅ on the support surface.

According to a third aspect, the present disclosure relates to a methodfor dehydrogenating an alkane to a corresponding olefin comprisingflowing the alkane through a reactor comprising a catalyst chamberloaded with the dehydrogenation catalyst of the present disclosure inany of its embodiments at a temperature in the range of 400-800° C. toform the corresponding olefin and a reduced catalyst.

The general nature of the alkane substrate is not viewed as particularlylimiting to the oxidative dehydrogenation described herein. As usedherein, “alkane” or “paraffin” unless otherwise specified refers to bothbranched and straight chain saturated primary, secondary and/or tertiaryhydrocarbons of typically C₁-C₁₀. It is equally envisaged that thepresent disclosure may be adapted to cycloalkanes referring to cyclizedalkanes containing one or more rings and substituted alkanes and/orsubstituted cycloalkanes referring to at least one hydrogen atom that isreplaced with a non-hydrogen group, provided that normal valencies aremaintained and that the substitution results in a stable compound. In apreferred embodiment, the alkane is at least one straight-chain linearalkane of C₁ to C₁₀, preferably C₂-C₆, more preferably C₂-C₄ selectedfrom the group consisting of ethane (C₂H₆), propane (C₃H₈), and a butane(C₄H₁₀, n-butane, isobutane) and the corresponding olefin is a lightolefin selected from the group consisting of ethylene, propylene, abutene (1-butene, (Z)-but-2-ene, (E)-but-2-ene, isobutylene(2-methylpropene)) and butadiene respectively, more preferably thealkane is ethane or propane and the corresponding olefin is ethylene orpropylene respectively, most preferably the alkane is propane and thecorresponding olefin is propylene. In certain embodiments, the alkanemay be sourced from other industrial processes such as those used in thepetrochemical industry. Feedstocks generated from petroleum including,but not limited to, ethane, propane, butane, naphtha, pet naphtha,pygas, light pygas, and gas oil may serve as substrates for the methodof dehydrogenating an alkane described herein. In some embodiments,these streams or feedstocks may be processed (i.e. hydroprocessed) priorto the dehydrogenation. In certain embodiments, the alkane may bepropane and the propane may be abundantly available from a natural gassource or a refinery off gas source.

As used herein, dehydrogenation refers to a chemical reaction thatinvolves the removal of hydrogen from a molecule. It is the reverseprocess of hydrogenation. The dehydrogenation reaction may be conductedon both industrial and laboratory scales. Essentially dehydrogenationconverts saturated materials to unsaturated materials anddehydrogenation processes are used extensively in fine chemicals,oleochemicals, petrochemicals and detergents industries. The mostrelevant industrial pathway in light olefin production is typicallysteam cracking; the alternative fluid catalytic cracking (FCC) is onlyable to produce desired olefins in small concentrations with significantcatalyst deactivation. The FCC catalytic dehydrogenation of alkanes ismore selective but the reaction characteristics pose inherentdifficulties and impose certain technical constraints. For example,thermal dehydrogenation is strongly endothermic and often requiresoperation at both high temperature and high alkane partial pressure. Theoxidative dehydrogenation (ODH) of an alkane, which couples theendothermic dehydrogenation of the alkane with the strongly exothermicoxidation of hydrogen avoids the need for excess internal heat input andconsumes hydrogen. The advantages of the alkane ODH reaction includethat the reaction is i) exothermic, ii) thermodynamically unrestricted,iii) operates at a much lower temperature, and iv) minimizes coke(CO_(x)) deposition ensuring long-term stability of the catalyst.

Under standard operating conditions, an alkane is converted to acorresponding olefin by oxidative dehydrogenation in the presence of thedehydrogenation catalyst described herein in accordance with thechemical equation represented by formula (II), wherein y is a positivewhole number, preferably y is 2, 3, or 4, more preferably y is 3 and thealkane converted is propane and the corresponding olefin is propylene.

$\begin{matrix}{{{C_{y}H_{{2y} + 2}} + {\frac{1}{2}V_{2}O_{5}}}->{{C_{y}H_{2y}} + {H_{2}O} + {\frac{1}{2}V_{2}O_{3}}}} & ({II})\end{matrix}$

In some embodiments the alkane to olefin conversion may be accompaniedby complete oxidation of the alkane or the olefin as side and/orsecondary reactions as represented in formula (III) and formula (IV),wherein y is a positive whole number, preferably y is 2, 3, or 4, morepreferably y is 3, and y is the sum of a and b (y=a+b). The yield ofalkenes or olefins obtained by oxidative dehydrogenation on catalysts islimited by alkene or alkane combustion to carbon oxides CO_(x) (i.e. COand CO₂). In some embodiments a=y and b=0 and CO₂ is the sole combustionproduct considered. The minimization of these undesirable consecutiveand/or parallel combustion reactions is a key in the development ofsuccessful oxidative dehydrogenation catalysts.

$\begin{matrix}{{{C_{y}H_{{2y} + 2}} + {\frac{1}{2}V_{2}O_{5}}}->{{a\;{CO}_{2}} + {b\;{CO}} + {\frac{\left( {{2y} + 2} \right)}{2}H_{2}O} + {\frac{1}{2}V_{2}O_{3}}}} & ({III}) \\{{{C_{y}H_{2y}} + {\frac{1}{2}V_{2}O_{5}}}->{{a\;{CO}_{2}} + {b\;{CO}} + {\frac{2y}{2}H_{2}O} + {\frac{1}{2}V_{2}O_{3}}}} & ({IV})\end{matrix}$

The performance of the oxidative dehydrogenation can be modulated byadjusting conditions including, but not limited to, temperature,pressure, reaction time and/or catalyst loading. One important objectivein developing oxidative dehydrogenation catalysts is to reduce thereaction temperature of the process to minimize energy consumption. In apreferred embodiment, the oxidative dehydrogenation of an alkane to acorresponding olefin is carried out a temperature in the range of400-800° C., preferably 450-750° C., preferably 500-700° C., preferably525-675° C., preferably 540-660° C., preferably 560-640° C., preferably580-620° C., or about 640° C. and preferably at approximately standardpressure (100 kPa, 1 bar, 14.5 psi, 0.9869 atm) such as for example10-20 psi, preferably 12-18 psi, preferably 14-16 psi, preferably14.25-15 psi, or approximately 14.4-14.8 psi. In a preferred embodiment,the catalyst-alkane feed contact time is in the range of 5-60 seconds,preferably 10-40 seconds, preferably 15-35 seconds, more preferably16-30 seconds, or about 17 seconds. In a preferred embodiment, thecatalyst loading or amount of catalyst present in the oxidativedehydrogenation reaction is in the range of 0.05-1.0 g of catalyst permL of alkane feed injected, preferably 0.10-0.80 g/mL, preferably0.15-0.60 g/mL, preferably 0.20-0.50 g/mL, preferably 0.25-0.45 g ofcatalyst per mL of alkane feed injected, or about 0.42 g/mL. Theconditions may vary from these ranges and still provide acceptableconditions for performing the oxidative dehydrogenation of an alkane toa corresponding olefin utilizing the dehydrogenation catalyst of thepresent disclosure.

Oxidative dehydrogenation catalysts are evaluated for their percentconversion of the alkane as well as their selectivity to a product (i.e.the corresponding olefin or CO_(x) (CO and/or CO₂). The definitions usedin calculating the conversion and selectivity are represented for themethod of the present disclosure using the oxidative dehydrogenationcatalyst are represented in formula (V) and formula (VI) respectively.

$\begin{matrix}{{{Alkane}\mspace{14mu}{conversion}};{{X_{alkane}(\%)} = {\frac{\sum_{j}{z_{j}n_{j}}}{{(y)n_{alkane}} + {\sum_{j}{z_{j}n_{j}}}} \times 100}}} & (V) \\{{{Selectivity}\mspace{14mu}{to}\mspace{14mu} a\mspace{14mu}{product}};{{S_{j}(\%)} = {\frac{z_{j}n_{j}}{\sum_{j}{z_{j}n_{j}}} \times 100}}} & ({VI})\end{matrix}$In these formulas, z_(j) and n_(j) are the number of atoms of carbon andmoles of gaseous carbon containing product j, respectively. The termn_(alkane) is the mole of unconverted alkane in the product stream (i.e.y=3 for propane, y=2 for ethane, etc.). Alternatively, the conversion ofalkane (%) can be thought of as moles of alkane converted divided bymoles of alkane fed multiplied by 100% and the selectivity to productcan be thought of as moles of product divided by the difference of molesof alkane reacted minus moles of product multiplied by 100%.

In one embodiment, the method of the present disclosure has an oxidativedehydrogenation alkane conversion rate as defined with formula (V) of upto 80%, preferably up to 70%, preferably up to 65%, preferably up to60%, preferably up to 55%, preferably up to 50%, preferably up to 45%,preferably up to 40%, preferably up to 35%, such as for example 10-80%,preferably 20-70%, preferably 25-65%, more preferably 30-60% and atleast 5%, preferably at least 10%, preferably at least 15%, preferablyat least 20%, preferably at least 25%. In another embodiment, the alkaneis ethane, propane, or butane and the method has an alkane conversion ofup to 80%, preferably up to 70%, preferably up to 65%, preferably up to60%, preferably up to 55%, preferably up to 50%, preferably up to 45%,preferably up to 40%, preferably up to 35%, such as for example 5-50%,preferably 10-45%, preferably 12-40%, more preferably 15-35%. In apreferred embodiment, the alkane is propane and the corresponding olefinis propylene and the method is performed with a catalyst-alkane feedcontact time or reaction time of 5-60 seconds, preferably 10-40 seconds,preferably 15-35 seconds, more preferably 16-30 seconds at a reactiondehydrogenation temperature of 500-700° C., preferably 525-675° C.,preferably 540-660° C., preferably 560-640° C. and the method has apropane conversion of 10-80%, preferably 20-70%, preferably 30-65%,preferably 40-60%.

In one embodiment, the method of the present disclosure has an oxidativedehydrogenation olefin selectivity relative to a total percentage ofproducts formed as defined with formula (VI) of at least 60%, preferablyat least 65%, preferably at least 70%, preferably at least 75%,preferably at least 80%, preferably at least 85%, preferably at least90%, preferably at least 95% such as for example 60-90%, preferably65-88%, preferably 70-86%, more preferably 75-85%. In anotherembodiment, the alkane is ethane, propane, or butane and the method hasan olefin selectivity relative to a total percentage of products formedof at least 60%, preferably at least 65%, preferably at least 70%,preferably at least 75%, preferably at least 80%, preferably at least85%, preferably at least 90%, preferably at least 95% such as forexample 60-90%, preferably 65-89%, preferably 70-88%, more preferably75-86%. In a preferred embodiment, the alkane is propane and thecorresponding olefin is propylene and the method is performed with acatalyst-alkane feed contact time or reaction time of 5-60 seconds,preferably 10-40 seconds, preferably 15-35 seconds, more preferably16-30 seconds at a reaction dehydrogenation temperature of 500-700° C.,preferably 525-675° C., preferably 540-660° C., preferably 560-640° C.and the method has a propylene selectivity relative to a totalpercentage of products formed of at least 60%, preferably at least 65%,preferably at least 70%, preferably at least 75%, preferably at least80%, preferably at least 85%, preferably at least 90%, preferably atleast 95% such as for example 60-95%, preferably 65-90%, preferably70-88%, more preferably 75-86%.

In a preferred embodiment, the method of the present disclosure isperformed with a catalyst-alkane feed contact time or reaction time of5-60 seconds, preferably 10-40 seconds, preferably 15-35 seconds, morepreferably 16-30 seconds at a reaction dehydrogenation temperature of500-700° C., preferably 525-675° C., preferably 540-660° C., preferably560-640° C. and the method has a CO₂ or complete combustion selectivityrelative to a total percentage of products formed that is less than theolefin selectivity, and the CO₂ selectivity is no more than 40%,preferably no more than 35%, preferably no more than 30%, preferably nomore than 25%, preferably no more than 20%, preferably no more than 15%,preferably no more than 10% such as for example 5-40%, preferably10-35%, preferably 20-30%. In a preferred embodiment, the method of thepresent disclosure has a CO selectivity relative to a total percentageof products formed that is less than the olefin selectivity and lessthan the CO₂ selectivity and the CO selectivity is no more than 38%,preferably no more than 35%, preferably no more than 30%, preferably nomore than 25%, preferably no more than 20%, preferably no more than 15%,preferably no more than 10% such as for example 5-40%, preferably10-35%, preferably 20-30%.

In a more preferred embodiment, the alkane is propane and thecorresponding olefin is propylene and the dehydrogenation catalyst ofthe present disclosure in any of its embodiments has a weight ratio ofcalcium oxide to alumina in the range of 1:0.5 to 1:4, preferably 1:0.75to 1:2, most preferably 1:1 and the method of the present disclosure isperformed with a catalyst-alkane feed contact time or reaction time of5-60 seconds, preferably 10-40 seconds, preferably 15-35 seconds, morepreferably 16-30 seconds at a reaction dehydrogenation temperature of500-700° C., preferably 525-675° C., preferably 540-660° C., preferably560-640° C. and the method has a propane conversion of at least 60%,preferably at least 62%, preferably at least 64%, preferably at least65%, preferably at least 66%, preferably at least 68%, preferably atleast 70%, preferably at least 75%, preferably at least 80%, preferablyat least 85%, preferably at least 90% and a propylene selectivityrelative to a total percentage of products formed of at least 80%,preferably at least 82%, preferably at least 84%, preferably at least85%, preferably at least 86%, preferably at least 88%, preferably atleast 90%, preferably at least 95%.

In a preferred embodiment, the method of the present disclosure andalkane oxidative dehydrogenation (ODH) reactions incorporating thedehydrogenation catalyst described herein are performed in a gas phaseoxygen-free environment or atmosphere. The presence of excess oxygeninside the reactor or catalyst chamber increases the combustion reactionand therefore CO_(x) production. Preferably, the amount of oxygenavailable for the reaction is controlled by the catalyst available, orlattice oxygen of the catalyst, specifically the vanadium oxide species.By this method, in reducing the catalyst loading or increasing thealkane feed to catalyst ratio one can further minimize the availableoxygen and decrease the combustion reaction, thus enhancing olefinselectivity.

In a preferred embodiment, the reactor is a fluidized bed reactor. Asused herein, a fluidized bed reactor (FBR) is a type of reactor devicethat can be used to carry out a variety of multiphase chemicalreactions. In this type of reactor, a fluid (gas or liquid) is passedthrough a granular solid material (usually a catalyst, preferablyspherically shaped) at high enough velocities to suspend the solid andcause it to behave as though it were a fluid. This process, known asfluidization, imparts many important advantages to the fluidized bedreactor. It is equally envisaged that the method of the presentdisclosure may be adapted to be performed in a fixed-bed reactor, butthis generally results in lower oxidative dehydrogenation catalystactivity.

The solid substrate (the catalytic material upon which the chemicalspecies react) material in a fluidized bed reactor is typicallysupported by a porous plate known as a distributor, distributor plate orsparger distributor. The fluid is then forced through the distributor upthrough the solid material. At lower fluid velocities, the solids remainin place as the fluid passes through the voids in the material. This isreferred to as a packed bed reactor. As the fluid velocity is increased,the reactor will reach a stage where the force of the fluid on thesolids is enough to balance the weight of the solid material. This stageis referred to as incipient fluidization and occurs at this minimumfluidization velocity. Once this minimum velocity is surpassed, thecontents of the reactor bed begin to expand and swirl around similar toan agitated tank or boiling pot of water. The reactor is now a fluidizedbed. Depending on the operating conditions and properties of the solidphase various flow regimes can be observed in this type of reactor.

The fluidized bed reactor technology has many advantages including, butnot limited to, uniform particle mixing, uniform temperature gradientsand the ability to operate the reactor in continuous state. Due to theintrinsic fluid-like behavior of the solid material, fluidized beds donot experience poor mixing as in packed beds. The complete mixing allowsfor a uniform product that can often be hard to achieve in otherreaction designs. The elimination of radial and axial concentrationgradients also allows for better fluid-solid contact, which is essentialfor reaction efficiency and quality. Many chemical reactions require theaddition or removal of heat. Local hot or cold spots within the reactionbed, often a problem in packed beds, are avoided in fluidized conditionssuch as the fluidized bed reactor. In other reactor types, these localtemperature differences, especially hot spots, can result in productdegradation. Thus fluidized bed reactors are well suited to exothermicreactions. The bed-to-surface heat transfer coefficients for fluidizedbed reactors are also high. The fluidized bed nature of these reactorsallows for the ability to continuously withdraw product and introducenew reactants into the reaction vessel. Operating at a continuousprocess state allows for the more efficient production and removesstartup conditions in batch processes.

In certain embodiments, the fluidizability, reactivity, and stability ofthe catalyst of the present disclosure or experimental laboratory scaleoxidative dehydrogenation reactions and/or reaction behaviors may bedemonstrated or evaluated in a Plexiglas unit with dimensions matchingthat of a CREC riser simulator. This type of reactor has a capacity of50-60 cm³, preferably 51-55 cm³ or about 53 cm³ and is a batch unitdesigned for catalyst evaluation and kinetic studies under fluidized bedreactor conditions. The major components of the CREC riser simulatorinclude, but are not limited to, a vacuum box, a series of samplingvalves, a timer, two pressure transducers and three temperaturecontrollers. The product gas may be analyzed by gas chromatography (GC)with a thermal conductivity detector (TCD) and flame ionization detector(FID).

The oxidative dehydrogenation method of the present disclosure may beperformed at various temperatures and contact times. In one embodiment,the contact times may be chosen to be consistent with catalyst reductiontemperature reported by temperature programmed reduction (TPR) analysis.In a typical procedure, the oxidized catalyst sample of the presentdisclosure is loaded into the reactor basket and the reactor basket ischecked for potential leaks. Following the leak test the system ispurged by flowing pure inert gas, preferably nitrogen or argon, mostpreferably argon. The temperature program is started to heat the reactorto the desired temperature. The inert gas flow is maintained to keep thereactor from any interference of gas phase oxygen. Once the reactorreaches a desired temperature, the inert gas flow is discontinued andthe reactor isolation valve is closed once a desired pressure level isreached. A vacuum pump may be used to evacuate the vacuum box down toless than 100 kPa, preferably less than 50 kPa, preferably less than 25kPa, preferably less than 20 kPa. In one embodiment, the catalyst may befluidized by rotating agitation, preferably by an impeller at a speed of100-5000 rpm, preferably 1000-4500 rpm, preferably 2000-4250 rpm,preferably 3000-4000 rpm. In another embodiment, no agitation (i.e. 0rpm) is necessary to fluidize the catalyst. The alkane feed is injectedinto the reactor using a preloaded gas tight syringe and the reactionproceeds for a pre-specified amount of time. At the termination point,the isolation valve between the reactor and vacuum box may automaticallyopen and transfer all reactant and products to the vacuum box foranalysis.

In a preferred embodiment, the method for the dehydrogenation of analkane to a corresponding olefin utilizing the dehydrogenation catalystof the present disclosure in any of its embodiments further comprises i)oxidizing at least a portion of the reduced catalyst in an oxygenenvironment separated from the catalyst chamber to regenerate thedehydrogenation catalyst of the present disclosure and ii) repeating theflowing and the oxidizing at least once with a less than 10% decrease inpercent conversion of the alkane, a less than 10% decrease inselectivity for the olefin relative to a total percentage of productsformed, or both. In this manner, the dehydrogenation catalyst can berecovered and reused in at least 2 reaction iterations, preferably atleast 3, preferably at least 4, preferably at least 5, preferably atleast 6, preferably at least 10, preferably at least 15, preferably atleast 20 reaction iterations.

The dehydrogenation catalyst of the present disclosure can be reformedor regenerated from the reduced catalyst; in this case the regenerationis the oxidation of the reduced vanadium species on the support surface.In a preferred embodiment, the regeneration is oxidation under air flowof the reduced catalyst and is performed at a temperature of up to 700°C., preferably up to 600° C., preferably up to 500° C., preferably up to400° C. for a period of time of up to 30 minutes, preferably up to 20minutes, preferably up to 15 minutes, preferably up to 10 minutes,preferably up to 5 minutes. In one embodiment, the reduced catalyst canflow out of the catalyst chamber to an additional chamber orre-oxidation chamber, be exposed to air flow to regenerate thedehydrogenation catalyst, and flow back to catalyst chamber for use insubsequent reaction iterations. In a preferred embodiment, catalystperformance remains stable in cycles in terms of alkane conversion andolefin selectivity indicating the catalyst's ability to be regeneratedwhich confirms catalyst stability at high temperatures. In a preferredembodiment, at least a portion of the reduced catalyst is oxidized toregenerate the dehydrogenation catalyst per reaction cycle, preferablyless than 60%, preferably less than 40%, preferably less than 30%,preferably less than 20%, preferably less than 15%, preferably less than10%, preferably less than 8% of the reduced catalyst is oxidized toregenerate the dehydrogenation catalyst per reaction cycle.

In a preferred embodiment, there is a less than a 10% change in percentalkane conversion between the first and second iteration, preferablyless than 5%, preferably less than 4%, preferably less than 3%,preferably less than 2%, preferably less than a 1% change in percentalkane conversion between the first and second iteration. In anotherembodiment, there is a less than a 20% change in percent alkaneconversion, preferably less than 15%, preferably less than 10%,preferably less than 5%, preferably less than a 2% change in percentalkane conversion between the first and twentieth iteration, preferablybetween the first and fifteenth iteration, preferably between the firstand tenth iteration, preferably between the first and fifth iteration,preferably between the first and fourth iteration, preferably betweenthe first and third iteration, preferably between the first and seconditeration.

In a preferred embodiment, there is a less than a 10% change in percentolefin selectivity relative to a total percentage of products formedbetween the first and second iteration, preferably less than 5%,preferably less than 4%, preferably less than 3%, preferably less than2%, preferably less than a 1% change in percent olefin selectivityrelative to a total percentage of products formed between the first andsecond iteration. In another embodiment, there is a less than a 20%change in percent olefin selectivity relative to a total percentage ofproducts formed, preferably less than 15%, preferably less than 10%,preferably less than 5%, preferably less than 2% change in percentolefin selectivity relative to a total percentage of products formedbetween the first and twentieth iteration, preferably between the firstand fifteenth iteration, preferably between the first and tenthiteration, preferably between the first and fifth iteration, preferablybetween the first and fourth iteration, preferably between the first andthird iteration, preferably between the first and second iteration.

The examples below are intended to further illustrate methods protocolsfor preparing and characterizing the dehydrogenation catalyst of thepresent disclosure. Further, they are intended to illustrate assessingthe properties and performance of these dehydrogenation catalysts. Theyare not intended to limit the scope of the claims.

Example 1

Catalyst Synthesis

The catalyst samples were prepared by an impregnation method throughsoaking with excess ethanol as solvent. The support material γ-Al₂O₃(surface area of 141 m²/g) was received from Inframat AdvancedMaterials, Manchester, UK, while CaO (surface area of 4 m²/g) wasreceived from Loba Chemie, India. Before metal loading, the γ-Al₂O₃ andCaO supports were calcined under pure N₂ flow at 500° C. for 4 hours toremove moisture and volatile compounds. The calcined γ-Al₂O₃ sample wasplace in a beaker and ethanol was added. A desired amount of vanadylacetyl acetonate and CaO were then added to the beaker and the mixturewas left under stirring for 12 hours. The mixture was then placed undersonication for 10 minutes. The mixture was filtered and dried at roomtemperature under atmospheric conditions evaporating ethanol. Followingthe natural evaporation and drying, the sample was placed in an oven at100° C. for 24 hours in order to slowly remove any remaining solvent.The dried sample was then reduced with hydrogen (10% H₂ and 90% Ar) at500° C. in a specially designed fluidized bed reactor. Finally, thereduced sample was calcined under air at 500° C. for 4 hours to obtainthe oxide form of the catalyst. After this treatment, the catalystdisplayed a yellow color indicating the presence of V₂O₅ on the supportsurface. In this manner, two catalyst samples were prepared with CaO toγ-Al₂O₃ weight ratios of 1:4 and 1:1, respectively while keepingconstant a 10 wt % vanadium loading (VO_(x)/CaO-γAl₂O₃ (1:4) andVO_(x)/CaO-γAl₂O₃ (1:1)). A third sample was prepared using pure CaO assupport and 10 wt % vanadium loading (VO_(x)/CaO). The measured BETsurface areas of the prepared catalysts ranged from 14 to 25 m²/g.

Example 2

X-Ray Diffraction (XRD) Analysis and Characterization of the PreparedCatalysts

The crystallographic structure of the catalyst samples and the baresupports were investigated using X-ray diffraction (XRD) analysis. TheXRD patterns of all the samples were recorded on a Rigaku Miniflexdiffractometer with monochromatic Cu Kα radiation of 1.5406×10⁻¹ nmwavelength, an electrical current of 50 mA, an electrical voltage of 10kV and a scan rate of 2° per minute (normal scan rate) within the 2θrange from 10°-90° with a 0.02 step size.

FIG. 1A shows the XRD patterns of the three VO_(x)/CaO-γ-Al₂O₃ catalystsamples (different CaO/γ-Al₂O₃ ratios) as well as bare CaO support, bareγ-Al₂O₃ support, and V₂O₅ for comparison. The XRD pattern of V₂O₅ showswell defined crystal structures at 2θ angles of 12.8°, 17.4°, 19.7°,24.1°, 28.2°, 43.3°, and 48.2°. The γ-Al₂O₃ samples give two peaks at 2θangles of 48° and 67°, which are consistent with previous studies. TheXRD pattern of CaO shows well defined reflections at 2θ angles of 32°,38° and 55°, this is also in line with the previously publishedliterature [R. Molinder, T. P. Comyn, N. Hondow, J. E. Parkerc, V.Duponta, In situ X-ray diffraction of CaO based CO₂ sorbents, EnergyEnviron. Sci., 5 (2012) 8958-8969.—incorporated herein by reference inits entirety]. FIG. 1B shows the XRD patterns of the three XRD patternsof the three VO_(x)/CaO-γ-Al₂O₃ catalyst samples with different amountsof CaO and γ-Al₂O₃ and with the same amount of VO_(N). For theVO_(x)/CaO-γAl₂O₃ (1:1) and VO_(x)/CaO-γAl₂O₃ (1:4) samples, the γ-Al₂O₃peaks appeared at 2θ angles of 48° and 67°. The peaks which appeared at32°, 38°, and 55° can be attributed to CaO. All three samples confirmedthese peaks and the intensity of the CaO peaks decreased when thecontent of CaO in the sample decreased. The 19.5° peak on theVO_(x)/CaO-γAl₂O₃ (1:1) and VO_(x)/CaO catalyst samples can be ascribedto V₂O₅ crystals. Small V₂O₅ peaks were also detected in the CaO sampleat 2θ angles above 60°. This observation indicates that the VOx speciesin the catalyst samples mainly appeared as a highly dispersed amorphousphase on the support samples.

The XRD patterns of the three catalyst samples show few peakscorresponding to the vanadium oxide species. The similar XRD patterns ofthe VO_(x)/CaO-γAl₂O₃ (1:4) and VO_(x)/CaO-γAl₂O₃ (1:1) samples furtherconfirmed the non-crystalline appearance of VO_(x) species. This can beascribed to the fact that the VO_(x) species in the catalyst samples hasa highly dispersed amorphous phase on the γ-Al₂O₃ and CaO surface. Thereis also the alternative possibility of the presence of small virtuallyundetectable by XRD V₂O₅ crystalline nanoparticles with a high level ofdispersion on the γ-Al₂O₃ and/or CaO support. This observation isconsistent with other available findings in previously publishedliterature. The other probable phases, the AlV₂O₉ and the CaV₂O₆ phasewere also not detected in any of the catalyst samples. One can inferfrom this observation that the reaction between the vanadium and thesupport materials γ-Al₂O₃ and/or CaO is negligible during the treatment,even at 750° C.

Example 3

Laser Raman Spectroscopy Analysis and Characterization of the PreparedCatalysts

The molecular structures of various metal oxide species supported onCaO-γ-Al₂O₃ and CaO were analyzed using a Horiba Raman spectrometerattached to a confocal microscope. For each experiment, 0.5 g of samplewas dehydrated under dry air for an hour at 500° C. and then cooled toambient temperature. Each sample was analyzed using a Raman spectrometerwith a thermoelectrically cooled CCD detector (−73° C.). An argon ionlaser line of 532 nm wavelength was used to excite the catalyst samples.The Raman spectrometer was used for measuring and recording the spectraproduced from the excitation with a resolution of 1 cm⁻¹ at roomtemperature.

FIG. 2A presents the Raman spectra of the samples that were obtained atambient temperature. FIG. 2B also includes the Raman spectra of bare CaOand γ-Al₂O₃ supports and V₂O₅ samples. The Raman spectra analysisindicates that all of the three catalyst samples VO_(x)/CaO,VO_(x)/CaO-γ-Al₂O₃ (1:1) and VO_(x)/CaO-γ-Al₂O₃ (1:4) contain bothmonovanadate and polyvanadate with minute crystal particles of V₂O₅. Thebroad bands in the range of 670-945 cm⁻¹ such as 870 cm⁻¹ are attributedto the stretching mode of the polyvanadate species (V—O—V). The 945-1030cm⁻¹ band is ascribed to the stretching mode of V═O. The narrow1030-1035 cm⁻¹ band and broad 1069 cm⁻¹ band are ascribed to thestretching mode of the V═O bond in isolated monovanadate surfacespecies. All other bands appearing around 100, 180, 235, 285, 325, 345,448, 520, 567, and 993 cm⁻¹ are ascribed to bulk V₂O₅ crystals. Inaddition, all the catalyst samples have slight peaks at 1030-1035 cm⁻¹which corresponds to monovanadate species.

Example 4

Fourier Transform Infrared (FTIR) Spectroscopy Analysis andCharacterization of the Prepared Catalysts

A Nicolet 6700 Thermo Fischer Scientific instrument was used to recordthe Fourier transform infrared (FTIR) spectra of the synthesizedcatalyst samples and the bare support γ-Al₂O₃ and CaO samples. Foranalysis, 3 mg of sample was uniformly mixed with 0.4 g of potassiumbromide. The infrared spectra of pelletized samples were later collectedin the range of 400-4000 cm⁻¹.

FIG. 3A displays the FTIR spectra of VO_(x)/CaO-γ-Al₂O₃ (1:4),VO_(x)/CaO-γ-Al₂O₃ (1:1), and VO_(x)/CaO catalysts as well as CaO,γ-Al₂O₃, and V₂O₅ samples for comparison. The strong infrared bands at3464 cm⁻¹, 1629 cm⁻¹, 880 cm⁻¹ and 821 cm⁻¹ as shown in the FTIR spectrarepresenting γ-Al₂O₃ are attributed to the stretching vibration of theAl—O bond [A. Imtiaz, M. A. Farrukh, M. Khaleeq-ur-rahman, R. Adnan,Micelle-Assisted Synthesis of Al₂O₃.CaO Nanocatalyst: Optical Propertiesand Their Applications in Photodegradation of 2,4,6-Trinitrophenol,Scientific World Journal. 2013 (2013) 1-11.—incorporated herein byreference in its entirety]. CaO has strong infrared bands correspondingto 450 cm⁻¹, 1410 cm⁻¹, and 3650 cm⁻¹ as shown in the spectra. Thesepeaks may be attributed to lattice vibrations of CaO [M. Sadeghi, M. H.Husseini, A Novel Method for the Synthesis of CaO Nanoparticle for theDecomposition of Sulfurous Pollutant, J. Appl. Chem. Res. 7 (2013)39-49.—incorporated herein by reference in its entirety]. Stronginfrared bands were observed at 833 cm⁻¹, 1014 cm⁻¹, and 1629 cm⁻¹ onthe V₂O₅ vanadium oxide curve.

FIG. 3B displays the FTIR spectra of VO_(x)/CaO-γ-Al₂O₃ (1:4),VO_(x)/CaO-γ-Al₂O₃ (1:1), and VO_(x)/CaO catalyst samples alone. Theabsorption peak at 450 cm⁻¹, 1410 cm⁻¹, and 3650 cm⁻¹ in all threecatalyst samples VO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), andVO_(x)/CaO confirms the presence of CaO in the catalysts. The band at1014 cm⁻¹ and 1629 cm⁻¹ confirms the presence of V₂O₅ in the catalystsamples and the band at 829 cm⁻¹ and 880 cm⁻¹ confirms the presence ofγ-Al₂O₃ in the VO_(x)/CaO-γ-Al₂O₃ (1:4) and VO_(x)/CaO-γ-Al₂O₃ (1:1)catalyst samples. The peak at 1014 cm⁻¹ corresponds to the strongterminal oxygen bond (V⁵⁺═O) [X. Zhou, G. Wu, J. Wu, H. Yang, J. Wangand G. Gao, Carbon black anchored vanadium oxide nanobelts and theirpost-sintering counterpart (V₂O₅ nanobelts) as high performance cathodematerials for lithium ion batteries, Phys. Chem. Chem. Phys, 16 (2014)3973-3982.—incorporated herein by reference in its entirety].

Example 5

Scanning Electron Microscopy (SEM) and Energy Dispersive X-RaySpectroscopy (EDXS) Analysis and Characterization of the PreparedCatalysts

The elemental analyses of the prepared samples were conducted usingenergy dispersive X-ray spectroscopy (EDXS). For analysis, the catalystsamples were dispersed on a stub that is tapped with copper. Each of thesamples were coated with gold in order to eliminate charge build up,obtain better contrast, and enhance visibility at magnification of up toone million times. The sample was analyzed by scanning electronmicroscopy (SEM), while ensuring that the microscope is aligned in orderto avoid a lack of sharpness and focus. An electron beam is incidentacross the catalyst sample resulting in the generation of secondary andback scattered electrons, which are used to form images and X-rays whichwere used to obtain elemental constitutions of the catalyst samples.

The SEM together with the EDXS analysis was carried out to determine themetal dispersion. FIG. 4 presents a representative field emissionscanning electron microscope image of one of the catalyst samples,VO_(x)/CaO-γ-Al₂O₃ (1:1). The images of the elemental distribution canbe used to envisage the quality of the dispersion. FIG. 5 shows thedispersion of the vanadium element over the oxygen carrier samples inthe VO_(x)/CaO-γ-Al₂O₃ (1:1) catalyst. It is evident that the vanadiumparticles are well dispersed on the CaO-γAl₂O₃ support. This indicatessuperior dispersion of the VO_(x) over the CaO-γAl₂O₃ support.

Example 6

Temperature Programmed Reduction-Oxidation (TPR/TPO) Characterization ofthe Prepared Catalysts' Reducibility and Oxygen Carrying Capacity

The reduction temperature and reducibility of the catalyst samples weredetermined using the temperature programmed reduction (TPR) technique. AMicrometrics AutoChem II 2920 analyzer was used to conduct H₂-TPRexperiments at 101.3 KPa. For TPR analysis, 0.05 g of catalyst samplewas loaded in a U-shaped quartz tube using glass wool to hold thecatalyst particles inside. The tube was inserted into retaining nuts andO-rings and then positioned in tube ports placed in a heater. Beforeanalysis the sample was pretreated under an argon (Ar) flow at 500° C.to remove any volatile components. After pretreatment, the sample wascompletely oxidized by circulating a gas mixture of 5% O₂ and helium(He) balance, at 500° C. with a heating rate of 10° C./min. The samplewas then cooled down to ambient temperature under argon flow to ensureflushing out of any gas phase O₂ that may have been trapped in thecatalyst bed. The temperature programmed reduction experiment wascarried out by circulating a gas stream of 10% H₂/Ar at 50 cm³/min. Atthese conditions, the sample temperature was raised from roomtemperature to 850° C. at a heating rate of 10° C./min. With theincreasing bed temperature, hydrogen begins to react with the solidphase metal oxides producing water vapor. This water vapor was trappedby circulating the exit stream through a clod trap containing molecularsieves. The water free outlet gas stream was passed through a calibratedthermal conductivity detector (TCD) which detects the variation of thehydrogen concentration due to the reduction of the catalyst samples.

TPR/TPO is an important technique for the characterization of gas phaseoxygen free ODH catalysts given that is simulates reduction/oxidation ofthe catalysts during the actual ODH reaction with propane. It givesinformation about the reducibility and regeneration ability of thecatalyst. Formula (VII) gives the equation of the TPR reaction andformula (VIII) gives the equation of the ODH of propane reaction.

$\begin{matrix}{{{V_{2}O_{5}} + {2H_{2}}}->{{V_{2}O_{3}} + {2H_{2}O}}} & ({VII}) \\{{{C_{3}H_{8}} + {\frac{1}{2}V_{2}O_{5}}}->{{C_{3}H_{6}} + {H_{2}O} + {\frac{1}{2}V_{2}O_{3}}}} & ({VIII})\end{matrix}$In can be seen that in both the TPR reaction (formula (VII)) and the ODHor propane reaction (formula (VIII)) V₂O₅ is reduced to V₂O₃. Incontrast, the TPO cycle represents the catalyst regeneration cyclefollowing the reduction in the TPR reaction. Formula (IX) gives theequation of the TPO reaction.V₂O₃+O₂→V₂O₅  (IX):In addition, the TPR/TPO data can be further processed to determine theoxygen carrying capacity of the catalysts for the oxidativedehydrogenation of propane without any additional gas phase oxygen(catalyst reduction cycle). Therefore, TPR analysis indicates thetemperature range of catalyst activation and the amount of availablelattice oxygen for the ODH of propane.

FIG. 6 presents the TPR profiles of the VO_(x)/CaO-γ-Al₂O₃ catalystswith different CaO to γ-Al₂O₃ ratios. It can be seen that all threecatalyst samples have similar reduction profiles and generally undergomost reduction between 350 and 620° C. The TPR profiles show that theVO_(x)/CaO catalyst sample has two humps between 95-287° C. and 300-430°C. while VO_(x)/CaO-γ-Al₂O₃ (1:1) shows only one hump between 260 and450° C. due to the highly reducible VO_(x) species that appeared on thesupport surfaces. The low temperature reduction hump with theVO_(x)/CaO-γ-Al₂O₃ (1:4) sample was less pronounced than the other twosamples. In addition to the initial reduction humps, all three catalystsamples exhibit a major reduction peak between 520-580° C. While theinitial reduction hump can be attributed to the reduction of bulk V₂O₅like surface species, the single major peak attributed to each catalystsample confirmed the presence of monomeric and polymeric VO_(x) speciesat the surface and the relative absence of crystalline V₂O₅nanoparticles, which gives the indication of high reducibility. For allthe catalyst samples, there was no peak attributed to CaO or Al₂O₃. Thisis due to the fact that calcium and aluminum are higher in theelectrochemical series as compared to vanadium and hydrogen. Indeed, thetemperatures that will be required for the reduction of CaO and Al₂O₃with hydrogen are higher than the temperature range considered in theTPR experiment. However, the reduction peak temperatures of the samplessignificantly varied with the variation of the CaO content in thecatalyst formulation. The peak temperature of the lowest CaO containingVO_(x)/CaO-γ-Al₂O₃ (1:4) sample was 515° C. With increasing CaO content,the VO_(x)/CaO-γ-Al₂O₃ (1:1) sample, the peak temperature shifted to560° C. The CaO supported VO_(x)/CaO sample shows the highest peaktemperature at 583° C. Previously, Bosc, et al. and Koranne, et al.reported similar reduction behavior for CaO containing vanadiumcatalysts [H. Bosc, J. K. Bert, J. G. Van Ommen, P. J. Gellings, Factorsinfluencing the temperature programmed reduction profiles of vanadiumpentoxide, J. Chem. Soc. Faraday Trans 80 (1984) 2479-2488; and M. M,Koranne, J. G. Goodwin, G. Marcelin, Characterization of silica andalumina supported vanadia catalysts using temperature programmedreduction, J. Catal. 148 (1994) 369-377.—each incorporated herein byreference in its entirety]. The shift of reduction temperature ispossibly due to the increased active site metal-support interactionintroduced by the addition of CaO

The TPR data was further processed to evaluate the degree of reductionfor the three catalyst samples. The degree of reduction can be definedas the percentage of VO_(x) reduced to the actual quantity of vanadiumoxide available in the catalyst. The exposed reducible VO_(x) wascalculated from the amount of hydrogen uptake evaluated using numericalintegration of the resulting temperature programmed reduction peak area.The mass of reducible vanadium oxide in the catalyst sample wasevaluated using molar volume of gas at standard temperature/pressure(STP), volume of hydrogen uptake, molecular weight of vanadium oxide andstoichiometric number of hydrogen in the gas-solid reaction involved inreduction. The percentage of vanadium oxide reduction can be calculatedusing the relation of formula (X) and formula (XI).

$\begin{matrix}{{\%\mspace{14mu}{reduced}} = {\frac{W_{V}}{W_{0\;}} \times 100}} & (X) \\{W_{V} = \frac{M\; W_{V} \times V_{H_{2}}}{v \times V_{g}}} & ({XI})\end{matrix}$In this formula, W_(V) is the amount of reduced vanadium (g), MW_(v) isthe molecular weight of vanadium (g/mol), V_(H2) is the volume ofreacted hydrogen (cm³ at STP), V_(g) is the molar volume of gas (cm³/molat STP), W₀ is the initial weight of vanadium (g) and ν is thestoichiometric number of hydrogen based on the reaction stoichiometrypresented in formula (I). Assuming that V₂O₅ is the initial reduciblecatalyst species on the support, then the reduction reaction equation offormula (VII) applies.V₂O₅+2H₂→V₂O₃+2H₂O  (VII):

Table 1 shows the hydrogen uptake and the percentage reduction of thecatalyst samples. It can be seen from this table that the hydrogenuptake was increased with increasing the CaO content in the catalystsamples. The higher hydrogen uptake is possibly due to the higherdispersion of vanadium species as observed in the low temperaturereduction humps of the VO_(x)/CaO-γ-Al₂O₃ (1:1) and VO_(x)/CaOcatalysts. The increased basicity with higher quantities of CaO alsocontributes to the increased hydrogen consumption, given the acidicnature of hydrogen gas, which has higher affinity to catalysts withhigher basicity. The hydrogen gas in solution is acidic and will havehigher reactivity with the catalyst of high basicity.

TABLE 1 Temperature programmed reduction (TPR) data comparing hydrogenconsumption for the three prepared catalyst samples of varying CaOcontent H₂ uptake (mmol/g) Uptake of (% reduction) % H₂ % Sample 1^(st)2^(nd) 3^(rd) 4^(th) Average Error (cm³ STP) reduction VO_(x)/CaO-γ- 1.91.8 1.8 1.8 1.8 1.6 2.12 48.21 Al₂O₃ (1:4) (48%) (46%) (45%) (45%)VO_(x)/CaO-γ- 2.6 2.5 2.4 2.4 2.5 2.5 2.87 65.27 Al₂O₃ (1:1) (65%) (63%)(62%) (62%) VO_(x)/CaO 2.9 2.8 2.8 2.7 2.8 1.8 3.18 72.32 (72%) (70%)(70%) (69%)

In order to assess the oxygen carrying capacity and stability underredox cycles, the catalyst samples were subjected to repeatedconsecutive reduction and re-oxidation TPR and TPO cycles. The hydrogenconsumption in each TPR cycle was measured using the calibrated TCDsignals. The percentage of reduction of the catalyst in each TPR cyclecalculated from the hydrogen uptake data is presented in Table 1. It isapparent that in the repeated TPR/TPO cycles, the hydrogen uptake forthe catalyst samples was within a 2.5% error range (Table 1). For allthree catalyst samples, the hydrogen uptake remains consistent over therepeated TPR/TPO cycles although the percentage reduction of each samplevaries as discussed above. This observation indicates that the oxygencarrying capacity of the catalyst remains stable over the repeated redox(TPR/TPO) cycles. Additionally, the stable value of the hydrogenconsumption indicates good stability of the present catalysts.

Example 7

NH₃-Temperature Programmed Desorption (NH₃-TPD) Characterization of thePrepared Catalysts' Acidity

The acidity and acid strength of the catalysts were investigated usingammonia in a temperature programmed desorption (TPD) analysis. TheNH₃-TPD desorption kinetic analysis also helps evaluating themetal-support interactions of the supported catalysts. In the context ofthe present disclosure, the NH₃-TPD experiments were conducted using anAutoChem II 2029 Analyzer received from Micromeritics, USA. Similar tothe TPR experiments, 0.05 g (or 0.5 g) of catalyst was first loaded intothe U-shaped quartz container and degassed for 2 hours at 500° C. underargon (Ar) flow at 30 mL/min. The sample was then cooled to 120° C. andbrought to saturation with ammonia using a NH₃/He gas mixture (5%NH₃/He) at a rate of 50 mL/min. Following the ammonia saturation, thesystem was purged with helium at a temperature of 100° C. at a rate of50 cm³/min in order to remove any gas phase ammonia in the system andunabsorbed ammonia trapped in the catalyst bed. For desorption analysis,the catalyst bed temperature was raised from room temperature to 750° C.at 10° C./min. The ammonia chemisorbed was desorbed as the temperatureelevated to 750° C. The ammonia concentration of the effluent gas wasmonitored by the thermal conductivity detector.

The acid sites of the three catalyst samples were characterized by TPDusing NH₃ as the basic probe molecule. The area of the TPD curve peakgives acid amount while the position of the peak indicates the aciddistribution in the catalyst samples. Ammonia TPD can distinguish sitesonly by sorption strength; hence its shortcoming lies in its inabilityto differentiate between Lewis and Bronsted acid sites. Ammonia was usedin this disclosure to make comparisons of the total acidity and acidstrength for catalyst samples with different CaO/Al₂O₃ ratios. FIG. 7shows the relationship between the desorption volume as a function ofthe temperature. It can be seen that all three samples show similar TPDprofiles although the peak intensity and desorption peaks are shiftedwith the variation of the CaO/Al₂O₃ ratios. The NH₃-TPD profile forVO_(x)/CaO-γ-Al₂O₃ (1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), and VO_(x)/CaOcatalyst samples showed an initial desorption peak at 183° C., 300° C.,and 302° C. respectively followed by a high temperature desorption peakat 676° C., 636° C., and 620° C. respectively. The intensity of the hightemperature desorption peaks were significantly higher than that of thelow temperature peaks. This indicates that the percentage of strong acidsites is much higher than the percentage of weak acid sites. The totalacidity of each catalyst sample was calculated by integrating thecalibrated TPD profiles. Table 2 shows the uptake of NH₃ by the threecatalyst samples and their respective temperatures of desorption. Thetotal acidity of the catalyst samples was decreased with the increasingof the CaO content due to the basic nature of the CaO.

TABLE 2 Catalyst acidity of the prepared catalyst samples of varying CaOcontent as measured by NH₃-temperature programmed desorption (NH₃-TPD)NH₃ Uptake Peak (mmol/g) Temperatures Low High (° C.) Temp. Temp. LowHigh 0.05 g 0.05 g Sample Temp. Temp. 0.50 g 0.50 g TotalVO_(x)/CaO-γ-Al₂O₃ (1:4) 183 676 0.0133 0.0666 0.0799 0.55 2.69 3.24(17%) (83%) VO_(x)/CaO-γ-Al₂O₃ (1:1) 300 636 0.0163 0.0612 0.0775 0.542.04 2.58 (21%) (79%) VO_(x)/CaO 302 620 0.0190 0.0539 0.0729 0.55 1.582.13 (26%) (74%)

Example 8

NH₃-Temperature Programmed Desorption (NH₃-TPD) Kinetics Analysis of thePrepared Catalysts

Ammonia desorption kinetics were evaluated to determine the active sitemetal-support interaction of the catalyst samples. The activation energyof ammonia desorption and the pre-exponential factors were estimated bymodeling the NH₃-TPD experimental data of each catalyst sample.Cvetanovic and Amenonmiya described the desorption rate as a function oftemperature which is based upon the following assumptions: i)temperature (T) of desorption has a linear relationship with time (t),ii) the rate of desorption is of first order in coverage, iii) theconcentration of ammonia gas through the catalyst bed is uniform, iv)desorbed ammonia has zero feasibility for re-adsorption, and v) thecatalyst's surface is homogeneous for the NH₃ adsorption, which meansthe desorption constant (k_(d)=k_(d0) exp(−E/RT) is independent of thesurface coverage. Suitable experimental conditions were selected inorder to satisfy the assumptions in i) and iii). A high flow of ammoniagas through the catalyst bed was maintained in order to satisfy theassumption in iv). Unimolecular desorption of ammonia was assumed inorder to consider the assumption in ii). The ammonia desorption rate ata uniform first order energy of desorption can be evaluated using acomponent balance of desorbing NH₃ in accordance with formula (XII).

$\begin{matrix}{r_{d} = {{{- V_{m}}\frac{d\;\theta}{dt}} = {k_{do}\theta\;{\exp\left\lbrack {{- \frac{E}{R}}\left( {\frac{1}{T} - \frac{1}{T_{m}}} \right)} \right\rbrack}}}} & ({XII})\end{matrix}$In this formula, T_(m) is the centering temperature in ° C., V_(m) isthe volume of NH₃ adsorbed at saturated conditions in mL/g, V_(d) is thevolume of ammonia desorbed at different temperatures in mL/g, θ is thesurface coverage of the adsorbed species, E is the energy of ammoniadesorption in kJ/mol, and k_(do) is the pre-exponential factor in mL g⁻¹min⁻¹. Temperature (T) in a TPD experiment has a linear relationshipwith time (t) given by formula (XIII), formula (XIV), formula (XV),formula (XVI), formula (XVII), and formula (XVIII) wherein T is thedesorption temperature at time (t).

$\begin{matrix}{T = {T_{0} + {\alpha\; t}}} & ({XIII}) \\{\frac{dT}{dt} = \alpha} & ({XIV}) \\{\frac{d\;\theta}{dt} = {{\frac{d\;\theta}{dT}\frac{dT}{dt}} = {\alpha\;\frac{d\;\theta}{dT}}}} & ({XV}) \\{\frac{d\;\theta}{dT} = {{- \frac{k_{do}}{\alpha\; V_{m}}}\theta\;{\exp\left\lbrack {{- \frac{E}{R}}\left( {\frac{1}{T} - \frac{1}{T_{m}}} \right)} \right\rbrack}}} & ({XVI}) \\{\theta = {1 - \frac{V_{d}}{V_{m}}}} & ({XVII}) \\{\frac{{dV}_{d}}{dT} = {\frac{k_{do}}{\alpha}\left( {1 - \frac{V_{d}}{V_{m}}} \right){\exp\left\lbrack {{- \frac{E}{R}}\left( {\frac{1}{T} - \frac{1}{T_{m}}} \right)} \right\rbrack}}} & ({XVIII})\end{matrix}$The first order ordinary differential equation was solved using theseparation of variable method to obtain formula (XIX).

$\begin{matrix}{V_{d} = {V_{m}\left( {1 - {\exp\left\lbrack {{\ln\left( {1 - \frac{V_{o}}{V_{m}}} \right)} - {\frac{k_{d}{RT}^{2}}{E\;\alpha\; V_{m}}\left\{ {{\exp\frac{- E}{R}\left( {\frac{1}{T} - \frac{1}{T_{m}}} \right)} - {\exp\frac{- E}{R}\left( {\frac{1}{T_{0}} - \frac{1}{T_{m}}} \right)}} \right\}}} \right\rbrack}} \right)}} & ({XIX})\end{matrix}$In this formula, V₀ and T₀ are initial volume desorbed in mL/g and theinitial desorption temperature. In this formula, R is the universal gasconstant in kJ mol⁻¹ K⁻¹ and the heating rate (α) was taken as 10°C./min.

FIG. 8 shows that the obtained experimental TPD data and the proposedmodel have good agreement for all catalyst samples. This proves thevalidity of the proposed desorption model. The TPD data was fitted inthe resulting equation (formula (XIII)) using a non-linear regressionanalysis tool of MATLAB. Hence, the desorption energies and thepre-exponential factors of each catalyst sample were obtained. The normof the residuals and the coefficient of correlation were calculated foreach catalyst sample using MATLAB and MINITAB software at a 95%confidence limit. Table 3 reports the energy of desorption for the threesynthesized catalysts. Statistical properties such as the correlationcoefficient R², norm of residuals and 95% confidence intervals wereconsidered in the analysis. The values of R² and residual norms for allthree catalyst samples are close to 1 and 0, respectively, whichindicates that the proposed desorption model is applicable.

TABLE 3 Estimated parameter of ammonia-TPD kinetics of the preparedcatalyst samples of varying CaO content at 10° C./min K_(do) Norm of E(mL/g/ resid- V_(NH3) V_(NH3) Sample (kJ/mol) min) × 10⁵ uals × 10⁴(mmol/g) (mL/g) VO_(x)/CaO-γ-Al₂O₃ (1:4) 39.2 3.8 2.1 3.24 72.6VO_(x)/CaO-γ-Al₂O₃ (1:1) 74.8 1.3 6.4 2.58 57.7 VO_(x)/CaO 96.3 0.5 22.72.13 47.6

The values in table 3 show that as the loading of CaO is increased andthat of γ-Al₂O₃ is decreased, the energy of desorption increases. Thiscan be explained based on the amount of ammonia uptake for each of thecatalyst samples. The catalyst with the highest desorbed ammonia has thelowest desorption energy while the one with the lowest desorbed ammoniahas the highest desorption energy. A similar observation was describedby Ghamdi, et al. on γ-Al₂O₃ supported VO_(x) catalysts where a higherdesorption energy corresponds to a lower amount of NH₃ adsorbed from thecatalysts. The increase in the activation energy can also be linked tothe heterogeneity of the catalyst samples. The interaction between themixed support and the active site (VO_(x)) also plays a significant rolein the value of energy required for ammonia desorption and during thegas-solid reactions involved during the oxidative dehydrogenation ofpropane under the gas phase oxygen free conditions. Weak active sitemetal-support interactions will enable high dispersion of the activesite which will in turn lead to an increased availability of the latticeoxygen for the oxidative dehydrogenation (ODH) reaction and an easyreaction between VO_(x) and propane/propylene as opposed to strongactive site metal-support interactions. Hence a weaker activesite-support interaction will require a lower energy of desorption whichmeans the VO_(x)/CaO-γ-Al₂O₃ (1:4) catalyst has the weakest activesite-support interaction. However, the moderate active sitemetal-support interactions as shown by the VO_(x)/CaO-γ-Al₂O₃ (1:1)catalyst can be favorable to achieve higher propylene selectivity.

Example 9

Evaluation of the Prepared Catalysts in the Fluidized OxidativeDehydrogenation (ODH) of Propane

The gas phase oxygen free oxidative dehydrogenation (ODH) experimentswere conducted in a fluidized CREC Riser Simulator (CREC: ChemicalReactor Engineering Centre). The CREC Riser Simulator, a bench scalefluidized reactor (53 cm³) is very useful for catalyst evaluation andkinetic studies. It has several advantages including simulatingfluidized conditions of a riser/downer reactor even with a small amountof catalyst, minimal mass transfer limitations by using small sizedcatalyst particles, constant residence time distributions, andcontrolled isothermal conditions. The CREC Riser Simulator reactoroperates alongside different accessories which include temperaturecontrollers, a gas chromatograph (GC), a vacuum box, a main powerswitch, a water pressure indicator and a push button selector. Greaterdetails of the CREC Riser Simulator can be found in Al-Ghamdi, etal.—2012.

Propane ODH runs were carried out at different temperatures ranging from550° C. to 640° C. while reaction contact times were varied between10-31 seconds. The reaction temperatures were selected within thereduction temperature range of the catalysts as determined by the TPRanalysis, given that the solid catalyst are the only source of oxygen.The ODH of propane experiments were conducted using 0.5 g of catalyst.The oxidized catalyst sample was loaded into the catalyst basket locatedin the lower shell of the main reactor body of the CREC Riser Simulator.Following the catalyst loading, the system was pressurized up to 30 psiat room temperature to perform a leak test. A stable pressure readingunder closed conditions confirmed the absence of any leak. The reactoris then ready to be heated to the desired temperature. During theheating period, the system was maintained under argon (Ar) flow to keepthe reactor from any air interference. Once the reactor reached thedesired temperature level, the argon flow was stopped. Consequently, thereactor pressure started to decrease sharply. The four port valve of theCREC Riser Simulator was closed, as the reactor pressure approached oneatm (14.7 psi). Following the isolation of the reactor, the vacuum pumpwas turned on to evacuate the vacuum box down to 20.7 kPa (3.75 psi). Apreloaded syringe was used to inject 1.2 mL of feed (propane) into thereactor after setting the impeller in motion. The pressure transducerwas used to record the pressure profile of the reactor. At the end ofthe reaction period, the reactor contents were evacuated into the vacuumbox. The analysis of the gas product contained in the vacuum box wascarried out with the aid of an online GC equipped with three differentpacked columns. Two of these columns are the carbon-1000 and carbon-1004columns which are used for separating the hydrogen, oxygen, nitrogen,argon, carbon (IV) oxide, and carbon (II) oxide gases and which areserially connected with the thermal conductivity detector (TCD). A flameionization detector (FID) was utilized in detecting the hydrocarbonssuch as propane, propylene, ethane, ethylene, and methane after theywere separated with a Haye SepD column. The evaluation of catalystperformance was based on propane conversion, selectivity and yield.Formula (V) gives the definition used in calculating propane conversion.Formula (VI) gives the definition used in calculating selectivity to aproduct.

$\begin{matrix}{{X_{C_{3}H_{8}}(\%)} = {\frac{\sum_{j}{z_{j}n_{j}}}{{3n_{propane}} + {\sum_{j}{z_{j}n_{j}}}} \times 100}} & (V) \\{{S_{j}(\%)} = {\frac{z_{j}n_{j}}{\sum_{j}{z_{j}n_{j}}} \times 100}} & ({VI})\end{matrix}$In these formulas, z_(j) and n_(j) are the number of atoms of carbon andnumber of moles of gaseous carbon containing product j respectively. Inaddition, n_(propane) is the number of moles of unconverted propane inthe product stream.

The gas phase oxygen free oxidative dehydrogenation (ODH) of propaneexperiments were conducted in a fluidized CREC Riser Simulator usingpure propane (99.95% purity) as feed. Before performing the actualcatalytic ODH runs, thermal experiments (without any catalyst) wereconducted to confirm the contribution of any thermal conversion. Thehighest reaction temperature (640° C.) was selected for the thermalexperiments. The GC analysis of the thermal runs' products showed mainlyunconverted propane and a trace amount of ethane and methane most likelydue to thermal cracking of propane in the absence of catalyst.

In the catalytic runs, the reaction temperature was varied between 550and 640° C., while the reaction time was attuned from 10 to 31 seconds.The product analysis of the preliminary runs contains unreacted propane,propylene and carbon dioxide. Under the studied reaction conditions, nohydrogen was detected, indicating the absence of cracking and/ordehydrogenation. FIG. 9 shows the propane (C₃H₈) conversion andpropylene (C₃H₆) and CO₂ selectivity and their error limits in repeatedruns. The propane conversion and product selectivity in the experimentalrepeats are found to be within 2.5% error limits (FIG. 9). Mass balanceswere established for each of the three repeat of each individual run andthe mass balance closed consistently in excess of 95%. From the productanalysis, one can consider the following three possible reaction stepsduring the fluidized ODH or propane in the absence of gas phase oxygen.Formula (VIII) gives the balanced equation of the ODH of propane topropylene reaction. Formula (XX) gives the equation of the completeoxidation of propane reaction. Formula (XXI) gives the equation of thecomplete oxidation of propylene reaction.

$\begin{matrix}{{{2C_{3}H_{8}} + {V_{2}O_{5}}}->{{2C_{3}H_{6}} + {2H_{2}O} + {V_{2}O_{3}}}} & ({VIII}) \\{{{C_{3}H_{8}} + {5V_{2}O_{5}}}->{{3{CO}_{2}} + {4H_{2}O} + {5V_{2}O_{3}}}} & ({XX}) \\{{{C_{3}H_{6}} + {\frac{9}{2}V_{2}O_{5}}}->{{3{CO}_{2}} + {3H_{2}O} + {\frac{9}{2}V_{2}O_{3}}}} & ({XXI})\end{matrix}$Therefore, it is important to identify reaction conditions in order toachieve the high propylene yields and suppress the complete combustionreactions which produce CO₂. With the above factors in mind, experimentswere conducted under different conditions to demonstrate the effects ofcertain parameters on the propane conversion and product selectivityincluding, but not limited to, (i) the consecutive propane injectionwithout catalyst regeneration, (ii) reaction temperatures, and (iii)contact times.

The successive oxidative dehydrogenation of propane experiments withoutcatalyst regeneration were conducted to demonstrate the effects of thedegree of catalyst reduction on the propane conversion and productdistribution. To ensure the same reaction conditions, the reactor wasloaded with 0.5 g of catalyst and the temperature was maintained at 640°C. Furthermore, in each run the same 1.2 mL of propane was injected andthe reactions were allowed to proceed for a consistent 17 seconds. FIG.10 plots the propane conversion over the successive injection of propaneruns. FIG. 11 plots the propylene and carbon oxide product selectivityover the successive injection of propane runs. FIG. 10 demonstrates thatall three VO_(x)/CaO—γ-Al₂O₃ catalysts give their highest propaneconversion in the first injection, which gradually decreases in thefollowing successive propane injections. The availability of the oxygenat the catalyst surface mainly contributed to the high propaneconversion in the first injection. The appreciable levels of catalystactivity after all four successive injections can be attributed to thelattice oxygen availability in the catalyst matrix. In contrast, thediminishing trend of the propane conversion is due to the progressiveconsumption of the lattice oxygen in the catalysts. Among the threecatalysts, VO_(x)/CaO-γ-Al₂O₃ (1:1) displays the highest propaneconversion (51%) and propylene selectivity. This is consistent with itsmoderate acidity, moderate active site metal-support interactions andbalanced oxygen carrying capacity in comparison to the other twocatalysts as observed in the TPR analysis.

FIG. 11 demonstrates that the selectivity of both the desired propyleneand undesired carbon dioxide vary significantly during the successivepropane injection runs. In contrast to propane conversion, the firstinjection gives the lowest propylene selectivity and highest carbondioxide selectivity. This indicates that the surface oxygen favors thecomplete oxidation of propane/propylene and the production of carbondioxide. The propylene selectivity significantly increased in the secondinjection and subsequently the incremental increases became minimal inthe remaining runs although there is an increasing trend still evident.This variation in selectivity indicates that certain amounts of latticeoxygen are required to maximize selectivity to propylene and minimizeselectivity to carbon dioxide.

This observation is in line with the fact that selectivity to propylenein the oxidative dehydrogenation of propane over VO_(x) based catalystsis affected positively by the energy that binds the lattice oxygen withthe catalyst [S. A. Al-Ghamdi, M. M. Hossain, H. I. de Lasa, Kineticmodeling of ethane oxidative dehydrogenation over VOx/Al₂O₃ catalyst ina fluidized-bed riser simulator, Ind. Eng. Chem. Res. 52 (2013)5235-5244.—incorporated herein by reference in its entirety]. At higheroxidation states of the catalyst the binding energy of the latticeoxygen is low (low active site metal-support interaction), whicheventually leads to combustion of propane/propylene to carbon oxides.Furthermore, the surface oxygen atoms on the fresh or regeneratedcatalyst are loosely bonded with the catalysts, which easily react withpropane/propylene to produce carbon dioxide. In this case, a selectivecatalyst surface would be obtained only after the adsorbed oxygen of thebulk V₂O₅ like surface species has been consumed via the first propaneinjection. It was after the consumption of adsorbed oxygen through thefirst propane injection that higher selectivity at the catalyst surfacecould be obtained.

When compared, VO_(x)/CaO-γ-Al₂O₃ (1:1) shows significantly higherpropylene selectivity and much lower carbon dioxide selectivity thatthat of the VO_(x)/CaO catalyst. This catalyst shows up to 96% propyleneselectivity while the higher CaO containing catalysts produce up to 83%propylene selectivity. This can be attributed to the moderate level ofacidity of VO_(x)/CaO-γ-Al₂O₃ (1:1) as demonstrated by the NH₃-TPDresults. This observation is also consistent with the XRD and TPRresults. The proper balance of CaO/Al₂O₃ influences the VO_(x)dispersion forming more isolated non-crystalline VO_(x) species, whichfavors the propylene formation and suppresses the complete oxidation toCO₂. Furthermore, the increased V-support interaction with the CaOpromoted sample, as revealed by the TPD kinetics analysis, may alsoexplain the controlled ODH reaction between propane and the latticeoxygen of the catalyst, resulting in enhanced propylene selectivity.Works have been published on ODH selectivity as a function of theoxidation state of vanadium based catalysts [V. Balcaen, I. Sack, M.Olea, G. B. Marin, Transient kinetic modeling of the oxidativedehydrogenation of propane over a vanadia-based catalyst in the absenceof O₂, Appl. Catal. A: Gen. 371 (2009) 31-42; and O. S. Owen, M. C.Kung, H. Kung, The effect of oxide structure and cation reductionpotential of vanadates on the selective oxidative dehydrogenation ofbutane and propane, Catal. Lett. 12 (1992) 45-50; and Creaser D,Andersson B, Hudgins R R, Silverston P L, Transient kinetic analysis ofthe oxidative dehydrogenation of propane, J. Catal. 182 (1999)264-269.—each incorporated herein by reference in its entirety]. Thesepublished works were focused on ODH reactions that utilize successiveinjections of alkanes in the absence of gas phase oxygen and indicatethat high selectivity for alkenes in ODH reactions can be obtained atcertain lattice oxygen of the vanadium based catalyst. Lopex-Nitro, etal. found that the selectivity to propylene and butylene with respectiveusage of propane and butane as the feed could be strongly influenced bythe reducibility of the vanadium based catalyst. Balcaen, et al. alsoobserved the same tend of ODH or propane over a vanadium based catalyst[V. Balcaen, I. Sack, M. Olea, G. B. Marin, Transient kinetic modelingof the oxidative dehydrogenation of propane over a vanadia-basedcatalyst in the absence of O₂, Appl. Catal. A Gen. 371 (2009)31-42.—incorporated herein by reference in its entirety]. Studies byAl-Ghamdi, et al. on ethane ODH over γ-Alumina supported vanadiumcatalyst in the absence of gas phase oxygen is important for theselective conversion of alkane to alkene with the binding energy oflattice oxygen as the main driver of the reaction.

FIG. 12 presents propane conversion at different reaction temperaturesand a constant 17 second contact reaction time. FIG. 13 presents productselectivity for the desired propylene and undesired carbon dioxide atdifferent reaction temperatures and a constant 17 second contactreaction time. These runs are conducted using oxidized catalysts. Aftereach run the catalyst was re-oxidized by circulating air through thecatalyst bed. It can be seen that the VO_(x)/CaO-γ-Al₂O₃ (1:4) sampleshows very low conversion at 550° C., which is consistent with itshigher initial reduction temperature as observed in the TPR analysis(FIG. 6). The VO_(x)/CaO—γ-Al₂O₃ (1:4) sample is mainly reduced between520° C. and 580° C. Therefore, there is only a small fraction of latticeoxygen available for reaction at 550° C. In contrast, both theVO_(x)/CaO-γ-Al₂O₃ (1:1) and VO_(x)/CaO catalyst samples show somereduction at low temperatures, which may contribute to the higherpropane conversions at 550° C. when using these two samples. Propaneconversion increased with the increasing reaction temperature as thelattice oxygen of the catalyst activates at higher temperature (FIG. 6,TPR analysis). Interestingly, with increasing the reaction temperature,all the catalyst samples showed increased propylene selectivity anddecreased carbon dioxide selectivity (FIG. 13). The variation in thedegree of reduction of the catalyst with reaction temperatures waslikely responsible for the rise in the selectivity of propylene. Athigher temperatures, the degree of catalyst reduction increases (FIG. 6,TPR analysis) as a result of the lower binding energy of lattice oxygen.At such higher degrees of reduction of the catalysts, the selectivepathway toward ODH is preferred over that for combustion as observed inthe successive propane injection experimental runs. The good selectivityto propylene can also be attributed to the non-formation of largermolecules due to the interaction of the mixed support and the activesite of each catalyst, as detected by XRD. Among the three studiedcatalysts, VO_(x)/CaO-γ-Al₂O₃ (1:1) shows the highest propyleneselectivity. The carbon dioxide selectivity with this catalyst is alsolower than that of the VO_(x)/CaO catalyst while slightly higher thanthe VO_(x)/CaO-γ-Al₂O₃ (1:4) catalyst. The superior propyleneselectivity of the VO_(x)/CaO-γ-Al₂O₃ (1:1) catalyst can be attributedto the moderate level of acidity of VO_(x)/CaO-γ-Al₂O₃ (1:1) asdemonstrated by the NH₃-TPD results.

Propane ODH experiments were carried out at 10, 17, 24 and 31 seconds inorder to study the effect of reaction time on propane conversion,propylene selectivity and carbon dioxide selectivity at a temperature of640° C. FIG. 14 presents propane conversion at different reaction timesand a constant 640° C. reaction temperature. FIG. 15 presents productselectivity for the desired propylene and undesired carbon dioxide atdifferent reaction times and a constant 640° C. reaction temperature. Itis evident that propane conversion for all catalyst samples increaseswith the reaction time (FIG. 14). The propylene selectivity slightlyincreases from 10 to 17 seconds and after that it decreases withreaction contact time above 17 seconds. Conversely, the carbon dioxideselectivity slightly decreases from 10 to 17 seconds and increases from17 to 31 seconds (FIG. 15). The decrease of propylene selectivity andincrease of carbon dioxide selectivity at higher contact time are mainlydue to the consecutive oxidation of propylene and/or complete oxidationof propane to carbon dioxide.

Therefore, the higher contact times favor high propane conversions whilesome contact time favor high propylene selectivity and low carbon oxideselectivity. The good selectivity to propylene obtained from the threecatalyst samples can be attributed to the high proportion ofmonovanadate VO_(x) species which was detected from the laser Ramanspectroscopy result. Again, the VO_(x)/CaO-γ-Al₂O₃ (1:1) catalyst showsthe highest propane conversion and propylene selectivity and lowestcarbon oxide selectivity. This can be attributed to the moderate levelof acidity of the VO_(x)/CaO-γ-Al₂O₃ (1:1) catalyst as depicted in theNH₃-TPD results.

Thus, one can conclude that the performance of the VO_(x)/CaO-γ-Al₂O₃(1:4), VO_(x)/CaO-γ-Al₂O₃ (1:1), and VO_(x)/CaO catalyst samples isstrongly influenced by both reaction contact times and temperatures andalso catalyst regeneration. It can be inferred that successive feedinjections are the best for the ODH reaction; hence it is only oncompletion of the successive reaction cycles that catalyst could beregenerated. This can be applied industrially using a fluidized bedreactor that has reactor-generator compartments and has the ability totransfer only a small percentage of the catalyst to the regenerator.

FIG. 16 presents propylene and carbon dioxide selectivities as afunction of propane conversion at constant temperature. It is evidentthat with increasing propane conversion, propylene selectivity decreaseswhich is compensated for by an increasing CO₂ selectivity. All the threecatalyst samples show similar trends. This indicates that propylene isthe primary reaction product of propane while CO₂ comes from deepoxidation of propane as well as consecutive oxidations of propane andpropylene.

Table 4 compares the performance of the catalyst of the presentdisclosure VO_(x)/CaO-γ-Al₂O₃ (1:1) with the performance of the ODHcatalysts reported in the past literature [T. V. M. Rao, G. Deo, KineticParameter Analysis for Propane ODH: V₂O₅/Al₂O₃ and MoO₃/Al₂O₃ Catalysts,AIChE J. 53 (2007) 2432-2442.—incorporated herein by reference in itsentirety]. There is an appreciable comparison of the propyleneselectivity of the VO_(x)/CaO-γ-Al₂O₃ (1:1) catalyst with the othercatalysts as shown. This indicates that the ODH of propane with latticeoxygen (gas phase oxygen free conditions) is promising to enhance thepropylene selectivity even at higher propane conversion.

TABLE 4 Comparison of the performance of the prepared VO_(x)/CaO-γ-Al₂O₃(1:1) catalyst with that of other ODH catalysts previously reported inthe literature Propane Reactivity conversion % Highest temperature (athighest propylene Catalyst (° C.) selectivity) selectivity % Reference17.5% MoO/γ-Al₂O₃ 380 1.3 77 Rao, et al. 17.5% MoO/γ-Al₂O₃ 380 1.7 83Rao, et al. 5% VO_(x)/γ-Al₂O₃ 550 11.7 86 Al-Ghamdi, et al.VO_(x)/CaO-γ-Al₂O₃ (1:4) 525-600 65.1 85.2 Current Disclosure

In conclusion, FTIR, Raman and XRD detected V₂O₅, CaO, and γ-Al₂O₃species in the prepared and synthesized catalysts, the XRD showed a verysmall amount of crystalline VO_(x) phases, the remaining VO_(x) appearedas an amorphous phase which is favorable for selective oxidations. Theisolated VOx phases, as observed by FTIR spectroscopy and XRD, arefavorable towards higher propylene selectivity. SEM images and elementalmapping showed good vanadium oxide dispersion of the catalytic materialon the mixed CaO-γ-Al₂O₃ support. Repeated TPR/TPO experiments showedthe consistent reduction and reoxidation behavior of the preparedcatalysts as well as the increase in oxygen carrying capacity withincreasing CaO content. Furthermore, NH₃-TPD analysis revealed that theacidity of the catalysts were progressively decreased with increasingCaO content and the activation energy of ammonia desorption decreasedwith increasing amounts of CaO reflecting the increased active sitemetal-support interactions which may control the reaction of the latticeoxygen and favor propylene as a selective product in ODH reactions ofpropane. Under gas phase oxygen free conditions the oxidativedehydrogenation of propane in the presence of the catalysts favored theselective formation of propylene product and minimized the completeoxidations to CO_(x) products. In addition, a higher degree of catalystreductions gave more selective products. The prepared catalyst withintermediate acidity and moderate active site metal-support interactions(VO_(x)/CaO-γ-Al₂O₃ (1:1)) displayed the highest propylene selectivity(85%) at higher propane conversion (65%).

Thus, the foregoing discussion discloses and describes merely exemplaryembodiments of the present invention. As will be understood by thoseskilled in the art, the present invention may be embodied in otherspecific forms without departing from the spirit or essentialcharacteristics thereof. Accordingly, the disclosure of the presentinvention is intended to be illustrative, but not limiting of the scopeof the invention, as well as other claims. The disclosure, including anyreadily discernible variants of the teachings herein, defines, in part,the scope of the foregoing claim terminology such that no inventivesubject matter is dedicated to the public.

The invention claimed is:
 1. A dehydrogenation catalyst, comprising: asupport material comprising alumina modified by calcium oxide, wherein aweight ratio of calcium oxide to alumina is from 1:0.2 to 1:8; and acatalytic material comprising one or more vanadium oxides disposed onthe support material; wherein the dehydrogenation catalyst comprises5-20% of the one or more vanadium oxides by weight relative to the totalweight of the dehydrogenation catalyst.
 2. The dehydrogenation catalystof claim 1, wherein the weight ratio of calcium oxide to alumina is 1:1.3. The dehydrogenation catalyst of claim 1, wherein the one or morevanadium oxides form an amorphous phase on the surface of the supportmaterial.
 4. The dehydrogenation catalyst of claim 1, wherein the one ormore vanadium oxides are at least one selected from the group consistingof V₂O₅, VO₂, and V₂O₃.
 5. The dehydrogenation catalyst of claim 4,which comprises at least 50% of V₂O₅ by weight relative to the totalweight of the one or more vanadium oxides.
 6. The dehydrogenationcatalyst of claim 1, which has an average particle size in the range of20-160 μm.
 7. The dehydrogenation catalyst of claim 1, which has anapparent particle density in the range of 1-10 g/cm³.
 8. Thedehydrogenation catalyst of claim 1, which has a BET surface area in therange of 5-50 m²/g.
 9. The dehydrogenation catalyst of claim 1, which isfluidizable and has Class B powder properties in accordance with Geldartparticle classification.
 10. A method for producing the dehydrogenationcatalyst of claim 1, comprising: mixing alumina with calcium oxide and avanadyl coordination complex or salt in a solvent to form loadedcatalyst precursors; reducing the loaded catalyst precursors with H₂ gasto form reduced catalyst precursors; and oxidizing the reduced catalystprecursors with oxygen to form the dehydrogenation catalyst.
 11. Amethod for dehydrogenating an alkane to a corresponding olefincomprising flowing the alkane through a reactor comprising a catalystchamber loaded with the dehydrogenation catalyst of claim 1 at atemperature in the range of 400-800° C. to form the corresponding olefinand a reduced catalyst.
 12. The method of claim 11, wherein the reactoris a fluidized bed reactor and the dehydrogenating is performed in a gasphase oxygen free environment.
 13. The method of claim 11, wherein thealkane is propane and the corresponding olefin is propylene.
 14. Themethod of claim 11, further comprising: oxidizing at least a portion ofthe reduced catalyst in a gas phase oxygen environment separated fromthe catalyst chamber to regenerate the dehydrogenation catalyst; andrepeating the flowing and the oxidizing at least once with a less than10% decrease in percent conversion of the alkane, a less than 10%decrease in selectivity for the olefin relative to a total percentage ofproducts formed, or both.
 15. The method of claim 11, wherein thedehydrogenation catalyst is present at an amount in the range of0.05-1.0 g of catalyst per mL of alkane.
 16. The method of claim 11,wherein the alkane is propane and the method has a propane conversion of10-80% at a reaction time of 5-60 seconds and a temperature of 500-700°C.
 17. The method of claim 11, wherein the alkane is propane and themethod has a propylene selectivity of at least 60% relative to a totalpercentage of products formed at a reaction time of 5-60 seconds and atemperature of 500-700° C.
 18. The method of claim 11, wherein thealkane is propane and the method has a carbon dioxide selectivity of nomore than 40% relative to a total percentage of products formed at areaction time of 5-60 seconds and a temperature of 500-700° C.
 19. Themethod of claim 11, wherein the alkane is propane, the dehydrogenationcatalyst has a weight ratio of calcium oxide to alumina of 1:1, and themethod has a propane conversion of at least 60% and a propyleneselectivity of at least 80% relative to a total percentage of productsformed.